RU2457197C2 - Oxidation system having secondary reactor for side stream - Google Patents

Oxidation system having secondary reactor for side stream Download PDF

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RU2457197C2
RU2457197C2 RU2008138888/04A RU2008138888A RU2457197C2 RU 2457197 C2 RU2457197 C2 RU 2457197C2 RU 2008138888/04 A RU2008138888/04 A RU 2008138888/04A RU 2008138888 A RU2008138888 A RU 2008138888A RU 2457197 C2 RU2457197 C2 RU 2457197C2
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preferably
reaction zone
reactor
reaction medium
specified
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RU2008138888/04A
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RU2008138888A (en
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Томас Эрл ВУДРАФФ (US)
Томас Эрл ВУДРАФФ
Алан Джордж УАНДЕРЗ (US)
Алан Джордж УАНДЕРЗ
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Истман Кемикал Компани
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    • BPERFORMING OPERATIONS; TRANSPORTING
    • B01PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01JCHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J10/00Chemical processes in general for reacting liquid with gaseous media other than in the presence of solid particles, or apparatus specially adapted therefor
    • B01J10/002Chemical processes in general for reacting liquid with gaseous media other than in the presence of solid particles, or apparatus specially adapted therefor carried out in foam, aerosol or bubbles
    • BPERFORMING OPERATIONS; TRANSPORTING
    • B01PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01JCHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/18Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
    • B01J8/1818Feeding of the fluidising gas
    • BPERFORMING OPERATIONS; TRANSPORTING
    • B01PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01JCHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J8/00Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes
    • B01J8/18Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles
    • B01J8/20Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles with liquid as a fluidising medium
    • B01J8/22Chemical or physical processes in general, conducted in the presence of fluids and solid particles; Apparatus for such processes with fluidised particles with liquid as a fluidising medium gas being introduced into the liquid
    • CCHEMISTRY; METALLURGY
    • C07ORGANIC CHEMISTRY
    • C07CACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C51/00Preparation of carboxylic acids or their salts, halides or anhydrides
    • C07C51/16Preparation of carboxylic acids or their salts, halides or anhydrides by oxidation
    • C07C51/21Preparation of carboxylic acids or their salts, halides or anhydrides by oxidation with molecular oxygen
    • C07C51/255Preparation of carboxylic acids or their salts, halides or anhydrides by oxidation with molecular oxygen of compounds containing six-membered aromatic rings without ring-splitting
    • C07C51/265Preparation of carboxylic acids or their salts, halides or anhydrides by oxidation with molecular oxygen of compounds containing six-membered aromatic rings without ring-splitting having alkyl side chains which are oxidised to carboxyl groups
    • BPERFORMING OPERATIONS; TRANSPORTING
    • B01PHYSICAL OR CHEMICAL PROCESSES OR APPARATUS IN GENERAL
    • B01JCHEMICAL OR PHYSICAL PROCESSES, e.g. CATALYSIS OR COLLOID CHEMISTRY; THEIR RELEVANT APPARATUS
    • B01J2219/00Chemical, physical or physico-chemical processes in general; Their relevant apparatus
    • B01J2219/00002Chemical plants
    • B01J2219/00027Process aspects
    • B01J2219/0004Processes in series

Abstract

FIELD: chemistry.
SUBSTANCE: invention relates to an improved method of producing a composition of aromatic dicarboxylic acid, involving (a) oxidation of a multiphase reaction medium in a primary oxidation reactor to obtain a first suspension; (b) further oxidation of at least a portion of said first suspension in a secondary oxidation reactor which is of the bubble column type, wherein the method further involves feeding an aromatic compound into said primary oxidation reactor, where at least about 80 wt % of said aromatic compound fed into said primary oxidation reactor is oxidised therein, wherein head gases are moved from the top of the secondary oxidation reactor into the primary oxidation reactor. Disclosed are an optimised process and equipment for more efficient and cheaper liquid-phase oxidation. Such liquid-phase oxidation is carried out in a bubble column type reactor which ensures a highly efficient reaction at relatively low temperatures. When the oxidised compound is para-xylene and the oxidation reaction product is crude terephthalic acid (TPA), such a product, TPA, can e purified and extracted using cheaper methods than when TPA is obtained using the conventional high-temperature oxidation process.
EFFECT: improved method of producing a composition of aromatic dicarboxylic acid.
30 cl, 4 tbl, 31 dwg

Description

FIELD OF TECHNOLOGY

The present invention relates to a method for producing a polycarboxylic acid composition. One aspect of the present invention relates to the partial oxidation of an aromatic compound (e.g., para-xylene) to produce a crude aromatic dicarboxylic acid (e.g., crude terephthalic acid), which can then be purified and isolated. Another aspect of the invention relates to an improved reactor system that provides a more efficient and economical oxidation process.

BACKGROUND

Oxidation reactions are used in a number of existing industrial processes. For example, liquid phase oxidation is currently used to oxidize aldehydes to acids (e.g., propionic aldehyde to propionic acid), oxidize cyclohexane to adipic acid and oxidize alkyl aromatics to alcohols, acids or diacids. A particularly significant industrial oxidation process of the latter category (oxidation of alkyl aromatic compounds) is the liquid-phase catalytic oxidation of para-xylene to terephthalic acid. Terephthalic acid is an important compound that finds a number of uses. The main use of terephthalic acid is to use as a raw material in the production of polyethylene terephthalate (PET, PET). PET is a well-known plastic used in large quantities around the world for the manufacture of products such as bottles, fibers and packaging material.

In a typical liquid phase oxidation process, including partial oxidation of para-xylene to terephthalic acid, a liquid phase feed stream and a gas phase oxidizer stream are introduced into the reactor and a multiphase reaction medium is obtained in the reactor. The liquid-phase feed stream introduced into the reactor contains at least one oxidizable organic compound (e.g., para-xylene), while the gas-phase oxidizer stream contains molecular oxygen. At least part of the molecular oxygen introduced into the reactor in the form of a gas is dissolved in the liquid phase of the reaction medium, which ensures the availability of oxygen for the liquid-phase reaction. If the liquid phase of the multiphase reaction medium contains an insufficient concentration of molecular oxygen (for example, if some parts of the reaction medium are “oxygen depleted”), undesirable side reactions can lead to the formation of impurities and / or target reactions can be slowed down. If the liquid phase of the multiphase reaction medium contains too little oxidizable compound, the reaction rate may be unacceptably slow. In addition, if the liquid phase of the reaction medium contains an excess concentration of the oxidizable compound, additional undesirable side reactions can cause the formation of impurities.

Conventional liquid phase oxidation reactors are equipped with mixing means for mixing the multiphase reaction medium contained therein. Stirring of the reaction medium is provided in order to stimulate the dissolution of molecular oxygen in the liquid phase of the reaction medium, to maintain relatively uniform concentrations of dissolved oxygen in the liquid phase of the reaction medium, and to maintain relatively uniform concentrations of oxidizable organic compound in the liquid phase of the reaction medium.

Mixing of the reaction medium subjected to liquid-phase oxidation is often achieved by mechanical means of mixing in containers, such as, for example, continuous mixing reactors (LDCs, CSTR). Although LDCs can provide thorough mixing of the reaction medium, LDCs have several disadvantages. For example, LDCs have a relatively high capital cost because they require expensive engines, hydrostatic bearings and drive shafts and / or complex mixing mechanisms. In addition, the rotating and / or oscillating mechanical components of conventional LDCs require regular maintenance. Labor costs and shutdown times associated with such maintenance increase operating costs for LDCs. However, even with regular maintenance, the mechanical agitation systems used in LDCs are prone to mechanical breakdowns and may require replacement within a relatively short period of time.

Bubble column reactors provide an attractive alternative to LDCs and other mechanically agitated oxidation reactors. Bubble column reactors provide mixing of the reaction medium without the use of expensive and unreliable mechanical equipment. Bubble-column reactors typically include an elongated vertical reaction zone within which the reaction medium is located. Stirring of the reaction medium in the reaction zone is mainly due to the natural buoyancy of gas bubbles rising up through the liquid phase of the reaction medium. The mixing due to natural buoyancy created in the reactors as a bubble column reduces the capital and operating costs relative to reactors with mechanical stirring. In addition, the essentially absence of moving mechanical parts associated with bubble column reactors provides an oxidation system that is less prone to mechanical breakdowns than mechanically stirred reactors.

When the liquid-phase partial oxidation of para-xylene is carried out in conventional oxidation reactors (LDCs or in a bubble column), the product withdrawn from the reactor is usually a suspension containing crude terephthalic acid (STK, CTA) and mother liquor. STK has a relatively high level of impurities (for example, 4-carboxybesaldehyde, para-toluic acid, fluorenones and other colored compounds), which makes it unacceptable as a raw material for the production of PET. Consequently, STK produced in conventional oxidation reactors is typically subjected to a purification process that converts STK into purified terephthalic acid (OTC, PTA) suitable for the production of PET.

One of the typical purification processes for converting STK into OTC involves the following steps: (1) replacing the mother liquor of the suspension containing the STK with water; (2) heating the STK / water suspension to dissolve the STK in water; (3) catalytic hydrogenation of the STK / water solution to convert impurities into more desirable and / or readily separable compounds; (4) precipitating the obtained OTC from the hydrogenation solution using a plurality of crystallization steps; and (5) isolating the crystallized OTC from the remaining liquids. Although this type of conventional cleaning process is effective, it can be very expensive. Some factors contributing to the increase in the cost of conventional STK purification methods include, for example, the thermal energy required to stimulate the dissolution of STK in water, the catalyst required for hydrogenation, the hydrogen stream required for hydrogenation, the loss of yield caused by the hydrogenation of a certain amount of terephthalic acid , and many vessels required for multi-stage crystallization. Therefore, it would be desirable to develop an oxidation system capable of producing STK that could be purified without the need for heat-stimulated dissolution in water, hydrogenation and / or multi-stage crystallization.

OBJECTS OF THE INVENTION

Thus, it is an object of the present invention to provide a more efficient and economical liquid phase oxidation system.

Another objective of the invention is to develop a more efficient and economical reactor and process for the liquid phase catalytic partial oxidation of para-xylene to terephthalic acid.

Another objective of the present invention is to provide a bubble column reactor, which contributes to an improved reaction of liquid phase oxidation with reduced formation of impurities.

Another objective of the present invention is to develop a more efficient and economical system for the production of pure terephthalic acid (PTC, PTA) by liquid-phase oxidation of para-xylene to produce crude terephthalic acid (CTK) and then purification of CTK to CTK.

Another objective of the present invention is to provide a bubble column reactor for the oxidation of para-xylene and the production of STK, capable of being purified without the need for heating stimulated dissolution of STK in water, hydrogenation of dissolved STK and / or multi-stage crystallization of hydrogenated STK.

It should be noted that the scope of the present invention, which is defined in the attached claims, is not limited to methods and equipment that are able to realize all of the above objectives. Moreover, the scope of the claimed invention may cover a number of systems that do not achieve all or any of the above objectives. Other objectives and advantages of the present invention will be readily apparent to those skilled in the art upon consideration of the following detailed description and the accompanying drawings.

SUMMARY OF THE INVENTION

One of the embodiments of the present invention relates to a method for producing a polycarboxylic acid composition, and this method includes the following steps: (a) oxidizing a multiphase reaction medium in a primary oxidation reactor, resulting in a first suspension; and (b) conducting further oxidation of at least a portion of the first slurry in the secondary oxidation reactor, wherein the secondary oxidation reactor is a bubble column reactor.

Another embodiment of the present invention relates to a reactor system. The reactor system includes a primary oxidation reactor and a secondary oxidation reactor. The primary oxidation reactor defines the boundaries of the first inlet and the first outlet. The secondary oxidation reactor is a bubble column reactor that defines a second inlet and a second outlet. The first outlet is connected by transmitting a fluid flow to the second inlet.

BRIEF DESCRIPTION OF THE DRAWINGS

Preferred embodiments of the present invention are described in more detail below with reference to the accompanying drawings, where

FIG. 1 is a side view of an oxidation reactor made in accordance with one embodiment of the present invention, in particular, illustrating the introduction of feed streams, oxidizing agent and reflux into the reactor, the presence of a multiphase reaction medium in the reactor, and the removal of gas and slurry from the upper and lower parts reactor, respectively;

FIG. 2 is an enlarged side view in section of the lower part of the reactor as a bubble column obtained along line 2-2 of FIG. 3, in particular, illustrating the location and configuration of the oxidizer bubbler used to introduce the oxidant stream into the reactor;

FIG. 3 is a plan view of the oxidizer bubbler of FIG. 2, in particular, illustrating that there are no openings for supplying an oxidizer in the upper portion of the bubbler;

FIG. 4 is a bottom view of the oxidizer bubbler of FIG. 2, in particular, illustrating the configuration of the oxidant supply openings in the lower portion of the bubbler;

FIG. 5 is a cross-sectional side view of an oxidizer bubbler obtained along line 5-5 of FIG. 3, in particular, illustrating the orientation of the oxidant supply openings in the lower portion of the oxidizer bubbler;

FIG. 6 is an enlarged cross-sectional side view of the bottom of the reactor as a bubble column, in particular, illustrating a system for introducing a feed stream into a reactor at multiple locations separated by vertical gaps;

FIG. 7 is a plan view taken along line 7-7 of FIG. 6, in particular, illustrating how the feed system shown in FIG. 6, distributes the feed stream in a preferred radial feed zone (ZS, FZ) and more than one azimuthal quadrant (Q 1 , Q 2 , Q 3 , Q 4 );

FIG. 8 is a sectional plan view similar to FIG. 7, but illustrating alternative means for supplying a feed stream to the reactor using bayonet tubes, each of which has many small holes for the feed;

FIG. 9 is an isometric view of an alternative system for introducing a feed stream into the reaction zone at multiple locations separated by vertical gaps without the need for multiple cuts into the vessel, in particular illustrating that the feed distribution system can at least partially rely on an oxidizer bubbler;

FIG. 10 is a side view of a single-sided feed distribution system and an oxidizer bubbler shown in FIG. 9;

FIG. 11 is a sectional top view taken along line 11-11 of FIG. 10, and further illustrates a single-sided feed distribution system based on an oxidizer bubbler;

FIG. 12 is a side view of a bubble column reactor equipped with internal and external reaction vessels;

FIG. 13 is an enlarged sectional view of a reactor as a bubble column of FIG. 12, obtained along line 13-13, in particular, illustrating the relative orientation of the internal and external reaction vessels;

FIG. 14 is a side view of an alternative bubble column reactor equipped with internal and external reaction vessels, in particular illustrating that the internal reaction vessel has a step diameter;

FIG. 15 is a side view of a bubble column reactor equipped with an external secondary oxidation reactor that receives suspension from a side fraction in the primary oxidation reactor;

FIG. 16 is a side view of a bubble column reactor equipped with a through-through external secondary oxidation reactor that receives a suspension from an enlarged hole in the side of the primary oxidation reactor;

FIG. 17a is a schematic side view of a reactor as a bubble column equipped with an internal structure for enhancing hydrodynamics in a reactor;

FIG. 17b is a sectional view of the reactor of FIG. 17a obtained along line 17b-17b of FIG. 17a;

FIG. 18a is a schematic side view of a bubble column reactor equipped with a first alternative internal structure to enhance reactor hydrodynamics;

FIG. 18b is a sectional view of the reactor of FIG. 18a obtained through line 18b-18b of FIG. 18a;

FIG. 19a is a schematic side view of a bubble column reactor equipped with a second alternative internal structure to enhance reactor hydrodynamics;

FIG. 19b is a sectional view of the reactor of FIG. 19a obtained through line 19b-19b of FIG. 19a;

FIG. 20a is a schematic side view of a bubble column reactor equipped with a third alternative internal structure to enhance reactor hydrodynamics;

FIG. 20b is a sectional view of the reactor of FIG. 20a obtained along line 20b-20b in FIG. 20a;

FIG. 21a is a schematic side view of a bubble column reactor equipped with a fourth alternative internal structure to enhance reactor hydrodynamics;

FIG. 21b is a sectional view of the reactor of FIG. 21a received on line 21b-21b in FIG. 21a;

FIG. 22a is a schematic side view of a bubble column reactor equipped with a fifth alternative internal structure to enhance reactor hydrodynamics;

FIG. 22b is a sectional view of the reactor of FIG. 22a obtained via line 22b-22b of FIG. 22a;

FIG. 23a is a schematic side view of a bubble column reactor equipped with a sixth alternative internal structure to enhance reactor hydrodynamics;

FIG. 23b is a sectional view of the reactor of FIG. 23a received along line 23b-23b of FIG. 23a;

FIG. 24a is a schematic side view of a bubble column reactor equipped with a seventh alternative internal structure to enhance reactor hydrodynamics;

FIG. 24b is a sectional view of the reactor of FIG. 24a received on line 24b-24b in FIG. 24a;

FIG. 25a is a schematic view of a bubble column reactor with a stepped diameter with hydrodynamic-enhancing internal structure;

FIG. 25b is a sectional view of the reactor of FIG. 25a obtained along line 25b-25b of FIG. 25a;

FIG. 26 is a side view of a bubble column reactor containing a multiphase reaction medium, in particular illustrating a reaction medium that is theoretically distributed across 30 horizontal thin layers of equal volume to quantify some gradients in the reaction medium;

FIG. 27 is a side view of a bubble column reactor containing a multiphase reaction medium, in particular illustrating first and second discrete 20 percent continuous volumes of the reaction medium that have substantially different oxygen concentrations and / or oxygen consumption rates;

FIG. 28A and 28B are enlarged views of crude terephthalic acid (STK) particles produced in accordance with one embodiment of the present invention, in particular illustrating that each STK particle is a low density particle with a high surface area consisting of a plurality of freely bonded STC subparticles;

FIG. 29A and 29B are enlarged views of a conventionally prepared CTK, particularly illustrating that a conventional CTK particle has an enlarged size, higher density, and lower surface area than the particle of the claimed STK of FIG. 28A and 28B;

FIG. 30 shows a simplified flow diagram of a prior art process for the production of purified terephthalic acid (OTC) and

FIG. 31 shows a simplified flow diagram of an OTC manufacturing process in accordance with one embodiment of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

One embodiment of the present invention relates to a liquid phase partial oxidation of an oxidizable compound. Such oxidation is preferably carried out in the liquid phase of a multiphase reaction medium in one or more stirred reactors. Suitable stirred reactors are, for example, bubble stirred reactors (e.g. bubble column reactors), mechanically stirred reactors (e.g. continuous mixing reactors), and flow stirred reactors (e.g. injection reactors). In one embodiment of the present invention, liquid phase oxidation is carried out using at least one bubble column reactor.

Used in this case, the definition of "bubble column reactor" will mean a reactor to facilitate chemical reactions in a multiphase reaction medium, where the mixing of the reaction mixture is mainly due to the upward movement of gas bubbles through the reaction medium. Used in this case, the definition of "mixing" will mean work distributed in the reaction medium, which causes fluid flows and / or mixing. The definitions used in this case, "majority", "predominantly" and "especially" will mean more than 50%. Used in this case, the definition of "mechanical mixing" will mean the mixing of the reaction medium, created by the physical movement of the rigid (them) or flexible (their) element (s) against or inside the reaction medium. For example, mechanical mixing can be achieved by rotation, oscillation and / or vibration of internal mixers, blades, vibrators or acoustic diaphragms located in the reaction medium. Used in this case, the definition of "flow mixing" will mean the mixing of the reaction medium caused by the high injection rate and / or recycling of one or more fluids in the reaction medium. For example, flow mixing may be provided by nozzles, ejectors, and / or eductors.

In a preferred embodiment of the present invention, less than about 40% of the mixing of the reaction medium in a bubble column reactor during oxidation is provided by mechanical stirring and / or flow mixing, more preferably less than about 20% of the mixing is provided by mechanical stirring and / or mixing with a stream, and most preferably less than about 5% of the mixing is provided by mechanical mixing and / or mixing with a stream. Preferably, the amount of mechanical stirring and / or stream stirring acting on the multiphase reaction medium during oxidation is less than about 3 kilowatts per cubic meter of reaction medium, more preferably less than about 2 kilowatts per cubic meter, and most preferably less than about 1 kilowatt per cubic meter.

In FIG. 1, a preferred bubble column reactor 20 is shown as consisting of a shell of a vessel 22, which has a reaction section 24 and a separation section 26. The reaction section 24 defines the boundaries of the reaction zone 28, while the separation section 26 defines the boundaries of the separation zone 30. Mostly a liquid-phase feed stream introduced into the reaction zone 28 through the inlet for raw materials 32A, b, c, d. Advantageously, a gas phase oxidant stream is introduced into the reaction zone 28 through an oxidizer bubbler 34 located at the bottom of the reaction zone 28. The liquid phase feed stream and the gas phase oxidizer stream together form a multiphase reaction medium 36 within the reaction zone 28. The multiphase reaction medium 36 contains a liquid phase and a gas phase. More preferably, the multiphase reaction medium 36 comprises a three-phase medium having solid phase, liquid phase and gas phase components. The solid-phase component of the reaction medium 36 preferably precipitates within the reaction zone 28 as a result of the oxidation reaction carried out in the liquid phase of the reaction medium 36. The bubble column reactor 20 includes a suspension outlet 38 located near the bottom of the reaction zone 28 and an outlet for gas 40, located near the top of the separation zone 30. The suspension effluent containing liquid-phase and solid-phase components of the reaction medium 36 is removed from the reaction zone 28 through the outlet an opening for slurry 38, while a predominantly gaseous effluent is discharged from the separation zone 30 through the gas outlet 40.

The liquid phase feed stream introduced into the reactor as a bubble column 20 through the feed inlet 32a, b, c, d, preferably contains an oxidizable compound, a solvent, and a catalyst system.

The oxidizable compound present in the liquid phase feed stream preferably contains at least one hydrocarbon group. More preferably, the oxidizable compound is an aromatic compound. Even more preferably, the oxidizable compound is an aromatic compound with at least one attached hydrocarbon group, or at least one attached substituted hydrocarbon group, or at least one attached heteroatom, or at least , with one attached carboxylic acid function (-COOH). Even more preferably, the oxidizable compound is an aromatic compound with at least one attached hydrocarbon group, or at least one attached substituted hydrocarbon group, each attached group containing from 1 to 5 carbon atoms. Even more preferably, the oxidizable compound is an aromatic compound containing precisely two attached groups, each attached group containing exactly one carbon atom and containing methyl groups and / or substituted methyl groups and / or at most one carboxylic acid group. Even more preferably, the oxidizable compound is para-xylene, meta-xylene, para-toluyl aldehyde, meta-toluyl aldehyde, para-toluic acid, meta-toluic acid and / or acetaldehyde. Most preferably, the oxidizable compound is para-xylene.

A “hydrocarbon group,” as defined herein, is at least one carbon atom that is bonded only to hydrogen atoms or to other carbon atoms. A “substituted hydrocarbon group,” as defined herein, is at least one carbon atom bonded to at least one heteroatom and at least one hydrogen atom. "Heteroatoms", as defined in this case, are all atoms other than carbon and hydrogen atoms. Aromatic compounds as defined herein contain an aromatic ring, preferably having at least 6 carbon atoms, even more preferably having only carbon atoms as part of an aromatic ring. Suitable examples of such aromatic rings include, but are not limited to, benzene, biphenyl, terphenyl, naphthalene and other carbon-based fused aromatic rings.

If the oxidizable compound present in the liquid phase feed stream is a conventional solid compound (that is, solid at standard temperature and standard pressure), it is preferable that the oxidizable compound be substantially dissolved in a solvent when introduced into the reaction zone 28. Preferably so that the boiling point of the oxidizable compound at atmospheric pressure is at least about 50 ° C. More preferably, the boiling point of the oxidizable compound is in the range of about 80 to 400 ° C, and most preferably in the range of 125 to 155 ° C. The amount of oxidizable compound present in the liquid phase feed stream is preferably in the range of about 2 to 40% by weight, more preferably in the range of about 4 to 20% by weight. and most preferably in the range of from about 6 to 15% of the mass.

It has now been noted that an oxidizable compound present in a liquid phase feed may include a combination of two or more different oxidizable chemicals. These two or more different chemical materials may be filed combined in a liquid phase feed stream or may be fed separately in multiple feed streams. For example, an oxidizable compound comprising para-xylene, meta-xylene, para-toluyl aldehyde, para-toluic acid and acetaldehyde may be fed into the reactor through a single inlet or through multiple separate inlets.

The solvent present in the liquid phase feed stream includes an acid component and an aqueous component. The solvent is preferably present in the liquid phase feed stream at a concentration in the range of about 60 to 98% by weight, more preferably in the range of about 80 to 96% by weight. and most preferably in the range from 85 to 94% of the mass. The acid component of the solvent is preferably a predominantly low molecular weight monocarboxylic acid containing 1-6 carbon atoms, more preferably 2 carbon atoms. Most preferably, the acid component of the solvent is predominantly acetic acid. Preferably, the acid component is at least up to about 75% by weight. solvent, more preferably at least about 80% of the mass. solvent and most preferably from 85 to 98% of the mass. solvent, the remainder being mainly water. The solvent introduced into the reactor as a bubble column 20 may include minor amounts of impurities, such as, for example, para-toluyl aldehyde, terephthalic aldehyde, 4-carboxybenzaldehyde (4-CBA, 4-CBA), benzoic acid, para-toluic acid , para-toluyl aldehyde, alpha-bromo-para-toluic acid, isophthalic acid, phthalic acid, trimellitic acid, polyaromatic compounds and / or suspended particulate materials. Preferably, the total amount of impurities in the solvent introduced into the reactor as a bubble column 20 is less than about 3% of the mass.

The catalyst system present in the liquid phase feed stream is preferably a homogeneous, liquid phase catalyst system capable of activating the oxidation (including partial oxidation) of the oxidizable compound. More preferably, the catalyst system comprises at least one polyvalent transition metal. Even more preferably, the polyvalent transition metal is cobalt. Even more preferably, the catalyst system includes cobalt and bromine. Most preferably, the catalyst system includes cobalt, bromine and manganese.

When cobalt is present in the catalyst system, it is preferable that the amount of cobalt present in the liquid phase feed stream is such that the cobalt concentration in the liquid phase of reaction medium 36 is maintained in the range of about 300 to 6,000 parts per million (ppm). ), more preferably in the range of from about 700 to 4200 parts by weight per million, and most preferably in the range of from 1200 to 3000 parts by weight per million. When bromine is present in the catalyst system, it is preferred that the amount of bromine present in the liquid phase feed stream is such that the concentration of bromine in the liquid phase of reaction medium 36 is maintained in the range of about 300 to about 5,000 ppm, more preferably in the range from about 600 to 4000 parts by weight per million, and most preferably in the range from 900 to 3000 parts by weight per million. When manganese is present in the catalyst system, it is preferable that the amount of manganese present in the liquid phase feed stream is such that the concentration of manganese in the liquid phase of reaction medium 36 is maintained in the range of about 20 to 1000 ppm, more preferably in the range of about from 40 to 500 parts per million, and most preferably in the range from 50 to 200 parts per million.

The concentrations of cobalt, bromine and / or manganese in the liquid phase of the reaction medium 36 above are expressed based on time averaged and volume averaged values. Used in this case, the term "time averaged" means the average value of at least 10 measurements obtained under the same conditions for a continuous period of time of at least 100 seconds. Used in this case, the concept of "averaged over the volume" means the average value of at least 10 measurements obtained in a homogeneous 3-dimensional space throughout the defined volume.

The mass ratio of cobalt to bromine (Co: Br) in the catalyst system introduced into reaction zone 28 is preferably in the range of about 0.25: 1 to 4: 1, more preferably in the range of about 0.5: 1 to 3: 1 and most preferably in the range of 0.75: 1 to 2: 1. The mass ratio of cobalt to manganese (Co: Mn) in the catalyst system introduced into reaction zone 28 is preferably in the range of about 0.3: 1 to 40: 1, more preferably in the range of about 5: 1 to 30: 1, and most preferably in the range of 10: 1 to 25: 1.

The liquid phase feed stream introduced into the reactor as a bubble column 20 may include small amounts of impurities, such as, for example, toluene, ethylbenzene, para-toluene aldehyde, terephthalic aldehyde, 4-carboxybenzaldehyde (4-CBA, 4-CBA), benzoic acid, para-toluic acid, para-toluic aldehyde, alpha-bromo-para-toluic acid, isophthalic acid, phthalic acid, trimellitic acid, polyaromatic compounds and / or suspended particulate materials. When a bubble column reactor 20 is used to produce terephthalic acid, meta-xylene and ortho-xylene are also considered impurities. Preferably, the total amount of impurities in the liquid-phase feed introduced into the reactor as a bubble column 20 is less than about 3% of the mass.

Although FIG. 1 illustrates an embodiment of the invention where the oxidizable compound, solvent and catalyst system are mixed together and introduced into the reactor as a bubble column 20 as a single feed stream, in an alternative embodiment of the invention, the oxidizable compound, solvent and catalyst system can be introduced separately into the reactor as a bubble column 20. For example, it is permissible to feed a para-xylene stream into the reactor as a bubble column 20 through an inlet separate from the inlet (s) for the solvent and catalyst.

Advantageously, the gas phase oxidant stream introduced into the reactor as a bubble column 20 through the oxidizer bubbler 34 contains molecular oxygen (O 2 ). Preferably, the oxidizing agent stream contains in the range of about 5 to 40 mol%. molecular oxygen, more preferably in the range of from about 15 to 30 mol%. molecular oxygen, and most preferably in the range from 18 to 24 mol%. molecular oxygen. Preferably, the remainder of the oxidizing agent stream is mainly gas or gases, such as nitrogen, which are inert to oxidation. More preferably, the oxidizing agent stream consists essentially of molecular oxygen and nitrogen. Most preferably, the oxidizing stream is dry air, which contains about 21 mol%. molecular oxygen and from about 78 to 81 mol%. nitrogen. In an alternative embodiment of the present invention, the oxidizing agent stream may contain substantially pure oxygen.

In FIG. 1, the bubble column reactor 20 is preferably equipped with a reflux distributor 42 located above the top surface 44 of the reaction medium 36. The reflux distributor 42 is operable to introduce droplets of a predominantly liquid phase reflux stream into the separation zone 30 using any droplet formation technique known in the art technicians. More preferably, the reflux dispenser 42 gives a droplet spray directed downward towards the upper surface 44 of the reaction medium 36. Preferably, this downward droplet spray has an effect (i.e., affects and affects) at least about 50% of the maximum horizontal transverse area section of the separation zone 30. More preferably, the spray of droplets affects at least 75% of the maximum horizontal cross-sectional area of the separation zone 30. Most preferably a good spray of droplets affects at least 90% of the maximum horizontal cross-sectional area of the separation zone 30. This downward spray of liquid reflux can help prevent the formation of foam on the upper surface or above the upper surface 44 of the reaction medium 36 and can also help to separate any drops of liquid or suspension trapped in a gas moving upward, which flows in the direction of the gas outlet 40. In addition, liquid reflux can serve as To reduce the amount of particulate material and the potential precipitation of compounds (e.g. dissolved benzoic acid, para-toluic acid, 4-CBA, terephthalic acid and metal catalytic salts) present in the gaseous effluent discharged from separation zone 30 through an outlet for gas 40. In addition, the introduction of droplets of reflux into the separation zone 30 due to distillation can be adapted to the composition of the gaseous effluent discharged through the gas outlet 40.

The liquid reflux stream introduced into the reactor as a bubble column 20 through a reflux distributor 42 preferably has approximately the same composition as the solvent component of the liquid phase feed stream introduced into the reactor as a bubble column 20 through a feed inlet 32a, b, c , d. Therefore, it is preferred that the liquid reflux stream contains an acid component and water. The acid component of the reflux stream is preferably low molecular weight monocarboxylic acid containing 1-6 carbon atoms, more preferably 2 carbon atoms. Most preferably, the acid component of the reflux stream is acetic acid. Preferably, the acid component is at least up to about 75% by weight. a reflux stream, more preferably at least about 80% of the mass. the flow of reflux and most preferably from 85 to 98% of the mass. reflux stream, the rest being water. Since the reflux stream generally has essentially the same composition as the solvent in the liquid phase feed stream, when reference is made to “all solvent” introduced into the reactor, such “all solvent” will include both the reflux stream and a solvent as part of the feed stream.

During liquid phase oxidation in a bubble column reactor 20, it is preferred that feedstocks, oxidizing agent and reflux streams are substantially continuously introduced into reaction zone 28, while effluent gas and slurry streams are substantially continuously withdrawn from reaction zone 28. Used in this case the term “substantially continuous” means a period of at least 10 hours, interrupted by less than 10 minutes. During oxidation, it is preferred that the oxidizable compound (e.g., para-xylene) is substantially continuously introduced into reaction zone 28 at a rate of at least about 8000 kilograms per hour, more preferably at a speed in the range of from about 15,000 to 200,000 kg / hour, even more preferably in the range of from about 22,000 to 150,000 kg / hour, and most preferably in the range of from about 30,000 to 100,000 kg / hour. Although it is generally preferred that the flow rates of the feed streams of the feed, oxidizing agent and reflux be substantially stationary, it should be noted that one embodiment of the present invention involves pulsating the flow of feeds of the feed, oxidizing agent and / or reflux in order to improve mixing and mass transfer. When the incoming stream of feed, oxidizing agent and / or phlegmy is pulsed, it is preferable that their costs range from about 0 to 500% of the stationary costs indicated in the present invention, more preferably from about 30 to 200% of the stationary the costs specified in the present invention, most preferably in the range from 80 to 120% of the stationary costs specified in the present invention.

The average volumetric reaction rate (SOS, STR) in a reactor as a bubble column 20 is defined as the mass of an oxidizable compound fed per unit volume of the reaction medium 36 per unit time (e.g., kilogram of para-xylene fed per cubic meter per hour). In normal use, the amount of the oxidizable compound not converted to the product, before calculating the SOS, it is usually necessary to subtract from the amount of the oxidizable compound in the feed stream. However, the conversion and yield are usually high for many oxidizable compounds preferred in this case (for example, para-xylene), and it is convenient to define this concept as described above. Due to capital costs and working materials, among others, it is generally preferable that the reaction be carried out with high SOS. However, carrying out a reaction with an increasingly high SOS can affect the amount or yield of partial oxidation. A bubble column reactor 20 is particularly useful when the SOS of an oxidizable compound (e.g., para-xylene) is in the range of about 25 to 400 kilograms per cubic meter per hour, more preferably in the range of about 30 to 250 kg / m 3 · Hour, even more preferably from about 35 to 150 kg / m 3 · hour and most preferably from 40 to 100 kg / m 3 · hour.

The oxygen SOS in the reactor as a bubble column 20 is defined as the mass of molecular oxygen consumed per unit volume of the reaction medium 36 per unit time (for example, a kilogram of molecular oxygen per cubic meter per hour). Due to capital costs and solvent consumption for oxidation, among others, it is generally preferable that the reaction be carried out with high oxygen SOS. However, carrying out the reaction with a higher and higher SOS of oxygen reduces the amount or yield of partial oxidation. Without reference to any theory, this is apparently related to the rate of transfer of molecular oxygen from the gas phase to the liquid phase at the surface area and from there to the volume of liquid. Too high SOC of oxygen is likely to lead to too low dissolved oxygen in the volume of the liquid phase of the reaction medium.

The generalized-average oxygen SOS is defined in this case as all the oxygen consumed in the entire volume of the reaction medium 36 per unit time (for example, a kilogram of molecular oxygen consumed per cubic meter per hour). A bubble column reactor 20 is particularly useful when the generalized-average oxygen SOS is in the range of about 25 to 400 kilograms per cubic meter per hour, more preferably in the range of about 30 to 250 kg / m 3 · hour, even more preferably about from 35 to 150 kg / m 3 · hour and most preferably from 40 to 100 kg / m 3 · hour.

During oxidation in a bubble column reactor 20, it is preferable that the ratio of the specific mass flow rate of the entire solvent (both from the feed stream and from the reflux stream) to the specific mass flow rate of the oxidizable compound entering the reaction zone 28 is maintained in the range of approximately from 2: 1 to 50: 1, more preferably in the range of about 5: 1 to 40: 1, and most preferably in the range of about 7.5: 1 to 25: 1. Preferably, the ratio of the specific mass flow rate of the total solvent introduced as part of the feed stream to the specific mass flow rate of the solvent introduced as part of the reflux stream is maintained in the range of about 0.5: 1 to the absence of any reflux stream, more preferably in the range from about 0.5: 1 to 4: 1, even more preferably in the range of from about 1: 1 to 2: 1, and most preferably in the range of from 1.25: 1 to 1.5: 1.

During liquid phase oxidation in a bubble column reactor 20, it is preferred that the oxidizing agent stream is introduced into the reactor as a bubble column 20 in an amount that provides molecular oxygen slightly above the stoichiometric oxygen demand. The amount of excess molecular oxygen required for the best results with a particular oxidizable compound has an effect on the overall economics of liquid phase oxidation. During liquid phase oxidation in a bubble column reactor 20, it is preferable that the ratio of the specific mass flow rate of the oxidizing agent to the specific mass flow rate of the oxidizable organic compound (e.g., para-xylene) entering the reactor 20 is maintained in the range of about 0.5: 1 to 20: 1, more preferably in the range of about 1: 1 to 10: 1, and more preferably in the range of 2: 1 to 6: 1.

In FIG. 1, the feed streams, oxidizing agent and reflux introduced into the reactor as a bubble column 20 together form at least part of a multiphase reaction medium 36. The reaction medium 36 is preferably a three-phase medium containing a solid phase, a liquid phase and a gas phase. As mentioned above, the oxidation of an oxidizable compound (eg, para-xylene) predominantly takes place in the liquid phase of the reaction medium 36. That is, the liquid phase of the reaction medium 36 contains dissolved oxygen and an oxidizable compound. The exothermic nature of the oxidation reaction that takes place in the reactor as a bubble column 20 causes a portion of the solvent (e.g., acetic acid and water) introduced through the feed inlet 32a, b, c, d to boil / evaporate. Therefore, the gas phase of the reaction medium 36 in the reactor 20 is formed mainly from the evaporated solvent and the undissolved, unreacted part of the oxidizing stream.

Some prior art oxidation reactors use heat transfer tubes / plates to heat or cool the reaction medium. However, such heat transfer structures may not be desirable in the inventive reactor and in the process described in this case. Therefore, it is preferable that the bubble column reactor 20 essentially does not include surfaces that come into contact with the reaction medium 36 and exhibits a time-averaged specific heat flux of more than 30,000 watts per square meter. In addition, it is preferable that less than about 50% of the time-averaged heat of reaction of the reaction medium 36 is removed using heat exchange surfaces, more preferably less than about 30% of the heat of reaction is removed using heat exchange surfaces, and most preferably less than about 10% of the heat of reaction is using heat transfer surfaces.

The concentration of dissolved oxygen in the liquid phase of the reaction medium 36 is a dynamic equilibrium between the mass transfer rate from the gas phase and the consumption rate during the reaction within the liquid phase (that is, it is not established simply due to the partial pressure of molecular oxygen in the supporting gas phase, although this and is one of the factors in replenishing the fraction of dissolved oxygen and really has an effect on the limiting upper concentration of dissolved oxygen). The amount of dissolved oxygen varies locally, being higher near the interface of the bubbles. In the general case, the amount of dissolved oxygen depends on the ratio of supply and consumption factors in various areas of the reaction medium 36. Over time, the amount of dissolved oxygen depends on the uniformity of gas and liquid mixing relative to the rates of chemical consumption. If you want to appropriately level the supply and consumption of dissolved oxygen in the liquid phase of the reaction medium 36, it is preferable that the time-averaged and volume-averaged oxygen concentration in the liquid phase of the reaction medium 36 is maintained above about 1 mol. ppm, more preferably in the range of from about 4 to 1000 mol. hours / million, even more preferably in the range of from about 8 to 500 mol. hours / million and most preferably in the range from 12 to 120 mol. ppm

The liquid phase oxidation reaction carried out in the reactor as a bubble column 20 is preferably a precipitation reaction that produces solids. More preferably, liquid phase oxidation carried out in a bubble column reactor 20 causes at least about 10% of the mass. an oxidizable compound (e.g., para-xylene) introduced into the reaction zone 28 to form a solid compound (e.g., crude terephthalic acid particles) in the reaction medium 36. Even more preferably, liquid phase oxidation causes at least about 50% of the mass. capable of oxidation of the compound to give a solid substance in the reaction medium 36. Most preferably, liquid phase oxidation causes at least about 90% of the mass. capable of oxidation of the compound to give a solid substance in the reaction medium 36. Preferably, the total amount of solids in the reaction medium 36 is more than about 3% of the mass. based on time averaging and volume averaging. More preferably, the total amount of solids in the reaction medium 36 is maintained in the range of about 5 to 40% by weight, even more preferably in the range of about 10 to 35% by weight. and most preferably in the range from 15 to 30% of the mass. Preferably, a substantial portion of the oxidation product (eg, terephthalic acid) produced in the bubble column reactor 20 is present in the reaction medium 36 as solids, as opposed to the remaining amount dissolved in the liquid phase of the reaction medium 36. The amount of solid phase of the product oxidation present in the reaction medium 36, preferably is at least about 25% of the mass. based on the entire oxidation product (solid and liquid phase) in the reaction medium 36, more preferably at least about 75% of the mass. based on the entire oxidation product in the reaction medium 36, and most preferably at least 95% of the mass. based on the entire oxidation product in the reaction medium 36. The numerical ranges given above for the amount of solids in the reaction medium 36 are essentially applicable for the stationary operation of the bubbler column 20 for a substantially continuous period of time, without starting, stopping or optimization operations a bubble column reactor 20. The amount of solids in the reaction medium 36 is determined using the gravimetric method. In the indicated gravimetric method, the corresponding part of the suspension is removed from the reaction medium and weighed. Under conditions that effectively maintain the overall solid-state distribution present in the reaction medium, free liquid is removed from a portion of the solids by precipitation or filtration, effectively without loss of precipitated solids, and with less than about 10% of the initial liquid mass remaining with a portion of the solids. The remaining liquid on solids is evaporated to dryness, effectively without sublimation of solids. The remainder of the solids is weighed. The ratio of the mass of a part of solids to the mass of the initial part of the suspension is the fraction of solids, usually expressed as a percentage.

The deposition reaction carried out in the reactor as a bubble column 20 can cause the formation of contamination (that is, the accumulation of solids) on the surface of some rigid structures that come into contact with the reaction medium 36. Therefore, in one embodiment of the present invention, it is preferred that the reactor as the bubble column 20, essentially did not include internal heat exchange, mixing or diverting the flow of structures in the reaction zone 28, since such structures will contribute to the image contamination. If internal structures are present in the reaction zone 28, it is desirable to exclude internal structures having external surfaces, which have a significant area facing upward flat surface, which is very prone to the formation of contaminants. Thus, if any internal structures are present in the reaction zone 28, it is preferred that less than about 20% of the total area of the upwardly exposed external surface of such internal structures forms substantially flat surfaces inclined less than about 15 degrees from the horizontal. Internal structures with this type of configuration are referred to as having a “non-contaminated” configuration.

As for FIG. 1, the physical configuration of the reactor as a bubble column 20 helps to ensure optimized oxidation of the oxidizable compound (e.g., para-xylene) with minimal formation of impurities. Preferably, the elongated reaction section 24 of the shell of the vessel 22 includes a substantially main cylindrical body 46 and a bottom 48. The upper end of the reaction zone 28 is defined by a horizontal plane 50 extending through the upper part of the main cylindrical body 46. The lower end 52 of the reaction zone 28 is determined by bottom inner surface of the bottom 48. Typically, the lower end 52 of the reaction zone 28 is located next to the outlet for the outlet for the suspension 38. Therefore, the elongated reaction it 28, defined inside the reactor as a bubble column 20, has a maximum length "L", measured from the upper end 50 to the lower end 52 of the reaction zone 28 along the axis of the elongated main cylindrical body 46. The length "L" of the reaction zone 28 is preferably in the range from about 10 to 100 meters, more preferably in the range of about 20 to 75 meters, and most preferably in the range of 25 to 50 meters. The reaction zone 28 has a maximum diameter (width) "D", which, as a rule, is equal to the maximum inner diameter of the main cylindrical body 46. The maximum diameter "D" of the reaction zone 28 is preferably in the range from about 1 to 12 meters, more preferably in the range from about 2 to 10 meters, even more preferably in the range of from about 3.1 to 9 meters, and most preferably in the range of from 4 to 8 meters. In a preferred embodiment of the present invention, reaction zone 28 has a length to diameter ratio of "L: D" in the range of about 6: 1 to 30: 1. Even more preferably, reaction zone 28 has an “L: D” ratio in the range of about 8: 1 to 20: 1. Most preferably, reaction zone 28 has an “L: D” ratio in the range of 9: 1 to 15: 1.

As discussed above, the reaction zone 28 of the bubble column reactor 20 adopts a multiphase medium 36. The reaction medium 36 has a lower end coinciding with the lower end 52 of the reaction zone 28 and an upper end located at the upper surface 44. The upper surface 44 of the reaction medium 36 defined along a horizontal plane that intersects the reaction zone 28 in a vertical position, where the contents of the reaction zone 28 passes from a continuous gas-phase state to a continuous liquid-phase state. The upper surface 44 is preferably located in a vertical arrangement, where the local time-averaged gas retention by a thin horizontal layer of the contents of the reaction zone 28 is 0.9.

The reaction medium 36 has a maximum height of "H", measured between its upper and lower ends. The maximum width “W” of the reaction medium 36 is typically equal to the maximum diameter “D” of the main cylindrical body 46. During the liquid phase oxidation in the reactor as a bubble column 20, it is preferred that the height “H” is maintained at about 60-120% of “ L ", more preferably from about 80 to 110% of" L ", most preferably from about 85 to 100% of" L ". In a preferred embodiment of the present invention, the reaction medium 36 has a height to width “H: W” ratio of greater than about 3: 1. More preferably, reaction medium 36 has an “H: W” ratio in the range of about 7: 1 to 25: 1. Even more preferably, reaction medium 36 has a height to width “H: W” ratio in the range of about 8: 1 to 20: 1. Most preferably, reaction medium 36 has a height to width "H: W" ratio in the range of about 9: 1 to 15: 1. In one embodiment of the invention, L = H and D = W so that the various sizes or ratios proposed for L and D are also applicable to H and W, and vice versa.

The relatively high L: D and H: W ratios proposed in accordance with an embodiment of the present invention can contribute to some significant advantages of the inventive system. As discussed in more detail below, it has been found that higher L: D and H: W ratios, as well as some other features discussed below, can induce favorable vertical gradients of molecular oxygen concentration and / or an oxidizable compound (e.g., para- xylene) in the reaction medium 36. In contrast to the generally accepted opinion, which favors a well mixed reaction medium with relatively uniform concentrations throughout, it has been found that vertical stepwise measurement concentrations of oxygen and / or a compound capable of oxidation contributes to a more effective and economical oxidation reaction. Minimizing the concentrations of oxygen and an oxidizable compound near the top of the reaction medium 36 can help to eliminate the loss of unreacted oxygen and an unreacted oxidizable compound through the upper gas outlet 40. However, if the concentrations of the oxidizable compound and unreacted oxygen are small throughout the reaction medium 36 , then the rate and / or selectivity of oxidation decreases. Therefore, it is preferable that the concentrations of molecular oxygen and / or oxidizable compound are significantly higher near the lower part of the reaction medium 36 than near the upper part of the reaction medium 36.

In addition, high L: D and H: W ratios cause the pressure in the lower part of the reaction medium 36 to be significantly higher than the pressure in the upper part of the reaction medium 36. This vertical pressure gradient is the result of the height and density of the reaction medium 36. One one of the advantages of a vertical pressure gradient is that increased pressure in the lower part of the vessel leads to greater oxygen solubility and greater mass transfer than would otherwise be achieved at comparable temperatures and head pressure in shallow reactors. Consequently, the oxidation reaction can be carried out at lower temperatures than would be expected in the case of less deep reactors. When a bubble column reactor 20 is used to partially oxidize para-xylene to crude terephthalic acid (STK), the ability to operate at lower temperatures with the same or better mass transfer of oxygen has several advantages. For example, the low temperature oxidation of para-xylene reduces the amount of solvent burned during the reaction. As discussed in more detail below, low temperature oxidation also contributes to the formation of small, high surface area, freely bonded, readily soluble STK particles, which can be subjected to more economical cleaning methods than large, low surface area, dense STK particles produced by conventional high temperature oxidation processes.

During oxidation in the reactor 20, it is preferable that the time-averaged and volume-averaged temperature of the reaction medium 36 is maintained in the range of about 125 to 200 ° C, more preferably in the range of about 140 to 180 ° C, and most preferably in the range of 150 to 170 ° C. Head pressure above reaction medium 36 is preferably maintained in the range of about 1 to 20 absolute bar (abs. Bar), more preferably in the range of about 2 to 12 abs. bar and most preferably in the range from 4 to 8 abs. bar. Preferably, the pressure difference between the top of the reaction medium 36 and the bottom of the reaction medium 36 is in the range of about 0.4 to 5 bar, more preferably the pressure difference is in the range of about 0.7 to 3 bar, and most preferably the pressure difference is from 1 to 2 bar. Although it is generally preferred that the head pressure above the reaction medium 36 is maintained at a relatively constant value, one embodiment of the present invention involves a pulsating head pressure to facilitate improved mixing and / or mass transfer in the reaction medium 36. When the head pressure pulsates, preferably so that the pulse pressure was in the range from about 60 to 140% of the head pressure in the stationary state specified in the present invention, more than dpochtitelno from about 85 to 115% of the head pressure in the steady state of the present invention, and most preferably from 95 to 105% of the head pressure in the steady state of the present invention.

Another advantage of the high “L: D” ratio of reaction zone 28 is that it can contribute to increasing the average reduced velocity of the reaction medium 36. The definition of “reduced velocity” and “reduced velocity of gas” used in this case with respect to the reaction medium 36, means the volumetric flow rate of the gas phase of the reaction medium 36 when rising in the reactor, divided by the horizontal cross-sectional area of the reactor at this elevation. The increased reduced speed provided by the high L: D ratio of the reaction zone 28 may contribute to local mixing and increased gas retention of the reaction medium 36. The time-averaged reduced velocities of the reaction medium 36 are at one fourth height, half height and / or three quarters of reaction height media 36 are preferably greater than about 0.3 meters per second, more preferably are in the range of about 0.8 to 5 m / s, even more preferably are in the range of from about 0.9 to 4 m / s, and most preferably in the range of from about 1 to 3 m / s.

In FIG. 1 section of the reactor compartment 26 as a bubble column 20 is simply an expanded portion of the shell of the vessel 22 located directly above the reaction section 24. The section of the reactor 26 reduces the speed of the upward flowing gas phase in the reactor as a bubble column 20 when the gas phase rises above the upper surface 44 of the reaction medium 36 and reaches the gas outlet 40. Such a decrease in the rate of rise of the gas phase facilitates the easy removal of trapped liquids and / or TV rdyh substances in the gas phase flowing upward and thus reduces undesirable loss of components present in the liquid phase of reaction medium 36.

The separation section 26 preferably in most cases includes a truncated cone-shaped transition wall 54, a cylindrical wide side wall 56 and an upper head part 58. The narrow lower end of the transition wall 54 is connected to the top of the main cylindrical body 46 of the reaction section 24. The wide upper end of the transition wall 54 connected to the bottom of the wide side wall 56. It is preferable that the transition wall 54 extends up and out from its narrow lower end at an angle in the range of from about 10 to 70 degrees from vert hiccups, more preferably in the range of about 15 to 50 degrees from the vertical, and most preferably in the range of about 15 to 45 degrees from the vertical. The wide side wall 56 has a maximum diameter "X", which is usually larger than the maximum diameter "D" of the reaction section 24, although when the upper part of the reaction section 24 has a smaller diameter than the entire maximum diameter of the reaction section 24, then X may be actually less than D. In a preferred embodiment of the present invention, the ratio of the diameter of the wide side wall 56 to the maximum diameter of the reaction section 24 "X: D" is in the range from about 0.8: 1 to 4: 1, most preferably in the range t 1.1: 1 to 2: 1. The upper head part 58 is connected to the top of the wide side wall 56. The upper head part 58 is preferably an elliptical head element defining a central opening that allows gas to leave the separation zone 30 through the gas outlet 40. On the other hand, the upper head part 58 can have any configuration, including conical. The separation zone 30 has a maximum height "Y", which is measured from the top 50 of the reaction zone 28 to the uppermost part of the separation zone 30. The ratio of the length of the reaction zone 28 to the height of the separation zone 30 "L: Y" is preferably in the range from about 2: 1 up to 24: 1, more preferably in the range of from about 3: 1 to 20: 1, and most preferably in the range of from 4: 1 to 16: 1.

As for FIG. 1-5, the location and configuration of the oxidizer bubbler 34 will be discussed in more detail. FIG. 2 and 3 show that the oxidizer bubbler 34 may include an annular element 60 and a pair of pipelines for the entry of oxidizer 64a, b. Typically, such pipelines for entering the oxidizing agent 64a, b can enter the vessel at a height above the ring element 60 and then turn down, as shown in FIG. 2. On the other hand, the oxidizing agent conduit may enter the vessel below the annular element 60 or approximately on the same horizontal plane as the annular element 60. Each oxidizing conduit 64a, b includes a first end connected to a corresponding inlet for an oxidizing agent 66a, b formed in the shell of the container 22, and a second end movably connected to the annular element 60. The annular element 60 is preferably formed from pipelines, more preferably from a plurality of sections Well pipelines, and most preferably from a plurality of sections of straight tubes rigidly attached to each other with the formation of a tubular polygonal ring. Preferably, the annular element 60 is formed of at least three straight pipe sections, more preferably from 6-10 pipe sections, and most preferably from 8 pipe sections. Accordingly, when the annular element 60 is formed of 8 tubular sections, it has an octagon configuration. Preferably, the pipe sections that form the oxidizing agent conduits 64a, b and the ring member 60 have a nominal diameter of greater than about 0.1 meters, more preferably in the range of about 0.2 to 2 m, and most preferably in the range of 0 25 to 1 m. As is apparently best shown in FIG. 3, it is preferable that there are essentially no holes formed in the upper part of the bubbler ring 60.

As is apparently best shown in FIGS. 4 and 5, the bottom of the oxidizer bubbler ring 60 has a plurality of openings for the oxidizer 68. The openings for the oxidizer 68 are preferably configured so that at least about 1% of the total throughput area the cross-section limited by the openings for the oxidizing agent 68 is located below the center line 64 (FIG. 5) of the annular element 60, where the center line 64 is located at the height of the volumetric center of gravity of the annular element 60. More preferably at least approximately 5% of the total cross-sectional area outlined by all the openings for the oxidizer 68 is located below the center line 64, with at least about 2% of the total cross-sectional area defined by the openings 68, which discharge the oxidizing stream predominantly downward within about 30 degrees from the vertical. Even more preferably, at least about 20% of the total cross-sectional area outlined by all of the oxidant openings 68 is located below the center line 64, wherein at least about 10% of the total cross-sectional area is defined by the openings 68 that discharge the oxidant stream mostly downward within about 30 degrees from the vertical. Most preferably, at least about 75% of the total cross-sectional area outlined by all the openings for the oxidizing agent 68 is located below the center line 64, with at least about 40% of the total cross-sectional area defined by the openings 68, which mainly discharge the oxidizing stream downward within approximately 30 degrees from the vertical. The fraction of the total cross-sectional area outlined by all of the oxidant openings 68 that are located above the center line 64 is preferably less than about 75%, more preferably less than about 50%, even more preferably less than about 25% and most preferably less than 5 %

As shown in FIG. 4 and 5, openings for oxidizing agent 68 include downward-opening holes 68a and beveled holes 68b. The downwardly directed openings 68a are configured to discharge the oxidant stream predominantly downward at an angle within approximately 30 degrees from the vertical, more preferably within approximately 15 degrees from the vertical, and most preferably within approximately 5 degrees vertically. As for FIG. 5, the beveled openings 68b are configured to discharge the oxidizing stream predominantly sideways and downward at an angle “A” that is in the range of about 15 to 75 degrees from the vertical, more preferably angle A is in the range of about 30 to 60 degrees from the vertical, and most preferably, angle A is in the range of 40 to 50 degrees from the vertical.

Preferably, substantially all of the openings for oxidizing agent 68 have approximately the same diameter. The diameter of the holes for oxidizing agent 68 is preferably in the range of about 2 to 300 millimeters, more preferably in the range of about 4 to 120 mm, and most preferably in the range of 8 to 60 mm. The total number of holes for oxidizing agent 68 in the annular element 60 is chosen so that it meets the criterion of low pressure drop, which is discussed in detail below. Preferably, the total number of holes for oxidizing agent 68 formed in the ring member 60 is at least about 10, more preferably the total number of holes for oxidizing agent 68 is in the range of about 20 to 200, and most preferably, the total number of holes for oxidizing agent 68 is in a range of from about 40 to about 100.

Although FIG. 1-5 illustrate a very specific configuration for an oxidizer bubbler 34, it should be noted that a number of oxidizer bubbler configurations can be used to achieve the advantages described in the present invention. For example, it is not necessary for the oxidizer bubbler to have the octagonal ring element configuration shown in FIG. 1-5. Moreover, it is permissible for the oxidizer bubbler to form any configuration of the flow line (s) in which a plurality of spaced openings are used to discharge the oxidant stream. The size, number and direction of outlet of the oxidizing holes in the flow conduit are preferably within the ranges presented above. In addition, the oxidizer bubbler is preferably configured to provide the azimuthal and radial distribution of molecular oxygen described above.

Regardless of the specific configuration of the oxidizer bubbler 34, it is preferable that the oxidizer bubbler be physically configured and oriented so as to minimize the pressure drop associated with the release of the oxidant stream from the flow line (s) through the oxidizer ports and into the reaction zone . Such a pressure differential is calculated as the time-averaged static pressure of the oxidizer flow inside the flow line at the outlet for the oxidizer 66a, b minus the time-averaged static pressure in the reaction zone at a height mark where one half of the oxidizer flow is introduced above this vertical location and the other half of the oxidant stream is introduced below this vertical location. In a preferred embodiment of the present invention, the time-averaged pressure drop associated with the release of the oxidant stream from the oxidizer bubbler is less than about 0.3 megapascal (MPa), more preferably less than about 0.2 MPa, even more preferably less than about 0, 1 MPa and most preferably less than 0.05 MPa.

Optionally, an oxidizer bubbler 34 may be provided with continuous or intermittent washing with a liquid (e.g., acetic acid, water and / or para-xylene) to prevent solid oxidation of the oxidizer bubbler. When using such a liquid wash, it is preferable that an effective amount of liquid (i.e., not only a small amount of liquid droplets, which may naturally be present in the oxidizer stream) pass through the oxidizer bubbler and exit from the oxidizer holes for at least one more period minutes every day. When the liquid is continuously or periodically discharged from the oxidizer bubbler 34, it is preferable that the time-averaged ratio of the specific mass flow rate of the liquid through the oxidizer bubbler to the specific mass flow rate of molecular oxygen through the oxidizer bubbler is in the range of about 0.05: 1 to 30: 1, or in the range of about 0.1: 1 to 2: 1, or even in the range of 0.2: 1 to 1: 1.

In many conventional bubble column reactors containing a multiphase reaction medium, substantially all of the reaction medium located below the oxidizer bubbler (or other mechanism for introducing the oxidizer stream into the reaction zone) has a very low gas retention value. As is known in the art, “gas retention” simply represents the volume fraction of a multiphase medium that is in a gaseous state. Zones of low gas retention in the medium can also be called "non-aerated" zones. In many conventional bubble column reactors, a significant portion of the total volume of the reaction medium is located below the oxidizer bubbler (or other mechanism for introducing the oxidizer stream into the reaction zone). Therefore, a significant part of the reaction medium located at the bottom of conventional reactors as a bubble column is non-aerated.

It was found that minimizing the number of non-aerated zones in the reaction medium subjected to oxidation in the reactor as a bubble column can minimize the formation of some undesirable impurities. Non-aerated zones of the reaction medium contain relatively few oxidizing bubbles. Such a low volume of oxidizer bubbles reduces the amount of molecular oxygen available for dissolution in the liquid phase of the reaction medium. Therefore, the liquid phase in the non-aerated zone of the reaction medium has a relatively low concentration of molecular oxygen. Such oxygen depleted, non-aerated zones of the reaction medium tend to stimulate undesirable side reactions rather than the target oxidation reaction. For example, when para-xylene is partially oxidized to form terephthalic acid, the insufficient availability of oxygen in the liquid phase of the reaction medium may cause the formation of undesirably high amounts of benzoic acid and conjugated aromatic rings, especially including highly undesirable colored molecules known as fluorenones and anthraquinones.

In accordance with one embodiment of the present invention, liquid phase oxidation is carried out in a bubble column reactor configured and oriented so that the volume fraction of the reaction medium with low gas retention values is minimized. Such minimization of non-aerated zones can be quantified by theoretical distribution of the entire volume of the reaction medium into 2000 discrete horizontal thin layers of the same volume. With the exception of the highest and lowest horizontal thin layers, each horizontal thin layer has a discrete volume bounded on its sides by the side wall of the reactor and bounded above and below by imaginary horizontal planes. The highest horizontal thin layer is bounded below by an imaginary horizontal plane, and above by the upper surface of the reaction medium. The lowest horizontal thin layer is bounded above by an imaginary horizontal plane, and below by the lower end of the tank. After a theoretical distribution of the reaction medium into 2000 discrete horizontal thin layers of equal volume, the time-averaged and volume-averaged gas retention of each horizontal thin layer can be determined. When such a method of determining the number of non-aerated zones is used, it is preferable that the number of horizontal thin layers having a time averaged and volume averaged gas retention of less than 0.1 is less than 30, more preferably less than 15, even more preferably less than 6 , even more preferably less than 4 and most preferably less than 2. Preferably, the number of horizontal thin layers having a gas retention of less than 0.2 is less than 80, more preferably less than 40, even more preferably less than 20, even more preferably less than 12 and most preferably less than 5. Preferably, the number of horizontal thin layers having a gas retention of less than 0.3 is less than 120, more preferably less than 80, even more preferably less than 40, even more preferably less than 20, and most preferably less than 15.

As for FIG. 1 and 2, it was found that the location of the oxidizer bubbler 34 lower in the reaction zone 28 provides several advantages, including reducing the number of non-aerated zones in the reaction medium 36. Given the height “H” of the reaction medium 36, the length “L” of the reaction zone 28 and a maximum diameter “D” of the reaction zone 28, it is preferred that a large portion (i.e.> 50% by weight) of the oxidizing agent stream be introduced into the reaction zone 28 within approximately 0.025H, 0.022L and / or 0.25D from the lower end 52 of reaction zone 28. More preferably, most of the sweat eye of the oxidizing agent is introduced into the reaction zone 28 in the range of about 0.02N, 0.018L and / or 0.2D from the lower end 52 of the reaction zone 28. Most preferably, the majority of the oxidant stream is introduced into the reaction zone 28 in the range of about 0.015N, 0.013L and / or 0.15D from the lower end 52 of reaction zone 28.

In the embodiment of FIG. 2, the vertical distance “Y 1 ” between the lower end 52 of the reaction zone 28 and the outlet of the upper openings for the oxidizer 68 of the oxidizer bubbler 34 is less than about 0.25 N, 0.022 L and / or 0, 25D, so that substantially the entire oxidant stream enters the reaction zone 28 within about 0.25H, 0.022L and / or 0.25D from the lower end 52 of the reaction zone 28. More preferably, Y 1 is less than 0.02H, 0.018L and / or 0.2D. Most preferably, Y 1 is less than 0.015 H, 0.013 L and / or 0.15 D, but more than 0.005 H, 0.004 L and / or 0.06 D. FIG. 2 illustrates a tangent line 72 at a location where the lower edge of the main cylindrical body 46 of the shell of the container 22 is connected to the upper edge of the elliptical lower head portion 48 of the shell of the tank 22. On the other hand, the lower head portion 48 may have any configuration, including a conical one, tangent line will also be determined by the lower edge of the main cylindrical body 46. The vertical distance «Y 2 'between the tangential line 72 and the top of oxidant sparger 34 preferably it is at least approximately mately 0,0012N, 0,001L and / or 0,001D; more preferably at least about 0.005 H, 0.004 L and / or 0.05 D and most preferably at least 0.01 H, 0.008 L and / or 0.1 D. The vertical distance “Y 3 ” between the lower end 52 of the reaction zone 28 and the outlet of the lower openings for the oxidizer 70 of the bubbler of the oxidizer 34 is preferably less than about 0.015 N, 0.013 L and / or 0.15 D; more preferably less than about 0.012H, 0.01L and / or 0.1D, and most preferably less than about 0.01H, 0.008L and / or 0.075D, but more than 0.003H, 0.002L and / or 0.025D.

In addition to the advantages provided by minimizing non-aerated zones (i.e., zones with low gas retention) in the reaction medium 36, it has been found that oxidation can be improved by maximizing the gas retention of the entire reaction medium 36. The reaction medium 36 preferably has a time-averaged and a volume-averaged gas retention of at least about 0.4, more preferably in the range of about 0.6 to 0.9, and most preferably in the range of 0.65 to 0.85. Some of the physical and technological characteristics of the bubble column reactor 20 contribute to the high gas retention discussed above. For example, in the case of a given reactor size and oxidizer flow rate, a high L: D ratio of reaction zone 28 gives a smaller diameter; this increases the reduced speed in the reaction medium 36, which, in turn, increases gas retention. In addition, the actual bubble column diameter and the L: D ratio are known to influence the average gas retention, even in the case of a given constant reduced velocity. In addition, minimization of non-aerated zones, especially in the lower part of reaction zone 28, contributes to an increased gas retention value. In addition, the head pressure and the mechanical configuration of the bubble column reactor can affect process stability at the high reduced speeds and gas retention values disclosed in the present invention.

As for FIG. 1, it has been found that an improved distribution of an oxidizable compound (e.g., para-xylene) in the reaction medium 36 can be achieved by introducing a liquid phase feed stream into the reaction zone 28 in a plurality of vertical-spaced locations. Preferably, the liquid phase feed stream is introduced into the reaction zone 28 through at least 3 feed holes, more preferably at least 4 feed holes. As used in this case, the definition of “openings for raw materials” means openings where a liquid-phase flow of raw materials is discharged into the reaction zone 28 for mixing with the reaction medium 36. It is preferred that at least 2 openings for the raw materials are separated by a vertical gap at least about 0.5D, more preferably at least about 1.5D, and most preferably at least about 3D. However, it is preferable that the highest opening for raw materials is separated from the lowest opening for raw materials by a vertical gap of not more than approximately 0.75N, 0.65L and / or 8D; more preferably not more than about 0.5H, 0.4L and / or 5D, and most preferably not more than about 0.4H, 0.35L and / or 4D.

Although it is desirable to introduce a liquid phase feed stream in a plurality of vertical locations, it has also been found to provide an improved distribution of the oxidizable compound in the reaction medium 36 if the bulk of the liquid phase feed stream is introduced into the lower half of the reaction medium 36 and / or reaction zone 28. Preferably, at least about 75% of the mass. a liquid phase feed stream is introduced into the lower half of the reaction medium 36 and / or reaction zone 28. Most preferably, at least 90% of the mass. a liquid phase feed stream is introduced into the lower half of the reaction medium 36 and / or the reaction zone 28. In addition, it is preferable that at least about 30% of the mass. a liquid phase feed stream was introduced into the reaction zone 28 within about 1.5 D from the lowest vertical location where the oxidizing stream is introduced into the reaction zone 28. This is the lowest vertical location where the oxidizing stream is introduced into the reaction zone 28, as a rule, down oxidizer bubbler; however, a number of alternative configurations for introducing an oxidizing agent stream into reaction zone 28 are contemplated by a preferred embodiment of the present invention. Preferably, at least about 50% of the mass. liquid phase feed is introduced within about 2.5 D from the lowest vertical location where the oxidizing agent stream is introduced into reaction zone 28. Preferably, at least about 75% of the mass. liquid phase feed is introduced within about 5D from the lowest vertical location, where an oxidizing stream is introduced into reaction zone 28.

Each raw material hole defines a cross-sectional area through which the raw material is discharged. Preferably, at least about 30% of the total cross-sectional area of all feed inlets is located within about 1.5 D from the lowest vertical location where the oxidant stream is introduced into reaction zone 28. Preferably, at least about 50% of the total cross-sectional area of all raw material inlets is within about 2.5 D of the lowest vertical location, where the oxidant stream is introduced into reaction zone 28. Pre respectfully, at least about 75% of the total opening area of all the feed inlets is located within about 5D of the lowest vertical location where the oxidant stream is introduced into reaction zone 28.

As for FIG. 1, in one embodiment of the present invention, the raw material inlets 32a, b, c, d are simply a series of vertically arranged openings along one of the walls of the container shell 22. Such raw material openings have substantially the same diameters that are less than about 7 centimeters, more preferably in the range of about 0.25 to 5 cm, and most preferably in the range of 0.4 to 2 cm. The bubble column reactor 20 is preferably equipped with a race control system ode liquid-phase feed stream out of each feed hole. Such flow control systems preferably include an individual choke 74a, b, c, d for each respective feed inlet 32a, b, c, d. In addition, it is preferable that the bubble column reactor 20 is equipped with a flow control system that allows at least a portion of the liquid-phase feed stream to be introduced into the reaction zone 28 at an increased reduced feed rate, which is at least about 2 meters per second, more preferably at least about 5 m / s, even more preferably at least about 6 m / s, and most preferably is in the range of from about 8 to 20 m / s. As used in this case, the definition of “reduced input rate” means the time-averaged volumetric flow rate of the feed stream from the feed hole, divided by the area of the feed hole. Preferably, at least about 50% of the mass. the feed stream is introduced into the reaction zone 28 at an increased reduced input rate. Most preferably, substantially the entire feed stream is introduced into the reaction zone 28 at an increased reduced feed rate.

In FIG. 6 and 7 show an alternative system for introducing a liquid phase feed stream into the reaction zone 28. In this embodiment, the feed stream is introduced into the reaction zone 28 at four different elevations. Each height mark is equipped with an appropriate raw material distribution system 76a, b, c, d. Each feed distribution system 76 includes a main feed pipe 78 and a collector 80. Each collector 80 is provided with at least two outlet openings 82, 84 connected to respective insertion pipes 86, 88, each of which extends into the reaction zone 28 of the vessel shell 22. Each insertion pipe 86, 88 provides a corresponding opening for the feedstock 87, 89 for discharging the feed stream into the reaction zone 28. The feedstock openings 87, 89 preferably have substantially the same diameters less than approximately about 7 cm, more preferably in the range of about 0.25 to 5 cm, and most preferably in the range of 0.4 to 2 cm. Preferably, the openings for raw materials 87, 89 of each raw material distribution system 76a, b, c, d are diametrically opposite each other so as to introduce a feed stream into the reaction zone 28 in opposite directions. In addition, it is preferable that the diametrically opposite openings for the raw materials 86, 88 of the adjacent distribution systems of raw materials 76 are oriented at 90 degrees of relative rotation of each other. During operation, the liquid-phase feed stream is launched into the main feed pipe 78 and then enters the collector 80. The collector 80 distributes the feed stream evenly for simultaneous introduction on opposite sides of the reactor 20 through the feed holes 87, 89.

FIG. 8 illustrates another configuration where each feed distribution system 76 is equipped with bayonet tubes 90, 92 rather than plug-in pipelines 86, 88 (shown in FIG. 7). The bayonet tubes 90, 92 extend into the reaction zone 28 and have many small holes for raw materials 94, 96 for discharging liquid phase raw materials into the reaction zone 28. Preferably, the small holes for raw materials 94, 96 of the bayonet tubes 90, 92 have substantially the same diameter of less than than about 50 mm, more preferably from about 2 to 25 mm, and most preferably from 4 to 15 mm.

FIG. 9-11 illustrate an alternative feed distribution system 100. Feed distribution system 100 introduces a liquid phase feed stream at a plurality of vertical and longitudinally spaced locations without the need for multiple taps into the side wall of the reactor as a bubble column 20. Raw material introduction system 100 typically includes a single inlet pipeline 102, manifold 104, a plurality of vertically mounted distribution tubes 106, a longitudinal support mechanism 108 and a vertical support mechanism m 110. The inlet pipe 102 penetrates the side wall of the main body 46 of the container shell 22. The inlet pipe 102 is movably connected to the collector 104. The collector 104 distributes the feed stream received from the inlet pipe 102 evenly across vertically mounted distribution pipes 106. Each distribution pipe 106 has a plurality of openings separated by vertical gaps for raw materials 112a, b, c, d for discharging a feed stream into the reaction zone 28. A longitudinal support mechanism 108 is connected to each distribution tube 106 and suppresses the relative longitudinal movement of the distribution tubes 106. The vertical support mechanism 110 is preferably connected to the longitudinal support mechanism 108 and to the top of the oxidizer bubbler 34. The vertical support mechanism 110 substantially suppresses the vertical movement of the distribution tubes 106 in the reaction zone 28. Preferably, the openings for Raw materials 112 had substantially the same diameters of less than about 50 mm, more preferably from about 2 to 25 mm, and most preferably from 4 about 15 mm. The vertical gaps for the holes for the raw materials 112 of the distribution system 110 shown in FIG. 9-11 may be substantially the same as described above for the feed distribution system of FIG. 1. Optionally, the openings for the raw materials may be elongated nozzles, rather than simple slots. Optionally, one or more devices for deflecting the flow may be located outside the flow line and on the trajectory of the liquids exiting from it into the reaction medium. Optionally, the opening near the bottom of the flow pipe may be sized to allow for solids to be cleaned within the liquid-phase feed distribution system, either continuously or intermittently. Optionally, mechanical devices such as hinged assemblies, check valves, overflow regulators, economizer valves, etc., can be used either to prevent solids from entering during operational disturbances, or to discharge accumulated solids from a liquid-phase feed distribution system.

It has been found that flow patterns of the reaction medium in many reactors as bubble columns can produce an uneven azimuthal distribution of the oxidizable compound in the reaction medium, especially when the oxidizable compound is introduced predominantly along one side of the reaction medium. Used in this case, the definition of "azimuthal" means the angle or gap about the vertical axis of the elongation of the reaction zone. Used in this case, the definition of "vertical" means within 45º from the vertical. In one embodiment of the present invention, a feed stream containing an oxidizable compound (e.g., para-xylene) is introduced into the reaction zone through a plurality of azimuthally spaced feed holes. Such raw material openings with azimuthal gaps can prevent regions with extremely high or extremely low concentrations of oxidizable compounds in the reaction medium. Various systems for introducing raw materials shown in FIG. 6-11 are examples of systems that provide a suitable azimuthal arrangement of holes for raw materials.

As for FIG. 7, in order to quantify the azimuthally distributed introduction of the liquid-phase feed stream into the reaction medium, the reaction medium can theoretically be distributed into four vertical azimuthal quadrants "Q 1 , Q 2 , Q 3 , Q 4 " of approximately the same volume. Such azimuthal quadrants “Q 1 , Q 2 , Q 3 , Q 4 ” are defined by a pair of imaginary intersecting perpendicular vertical planes “P 1 , P 2 ” extending above the maximum vertical size and maximum radial size of the reaction medium. When the reaction medium is in a cylindrical vessel, the line of intersection of the imaginary intersecting vertical planes P 1 , P 2 will approximately coincide with the vertical center line of the cylinder, and each of the azimuthal quadrants Q 1 , Q 2 , Q 3 , Q 4 will typically represent a wedge-shaped vertical volume having a height equal to the height of the reaction medium. Preferably, a significant portion of the oxidizable compound is discharged into the reaction medium through feed openings located in at least two different azimuthal quadrants.

In a preferred embodiment of the invention, not more than approximately 80% of the mass. capable of oxidation of the compound is discharged into the reaction medium through openings for raw materials, which can be located in one azimuthal quadrant. More preferably not more than 60% of the mass. capable of oxidation of the compound is discharged into the reaction medium through openings for raw materials, which can be located in one azimuthal quadrant. Most preferably not more than 40% of the mass. capable of oxidation of the compound is discharged into the reaction medium through openings for raw materials, which can be located in one azimuthal quadrant. Such azimuthal distribution parameters of the oxidizable compound are measured when the azimuthal quadrants are azimuthally oriented such that the maximum possible amount of the oxidizable compound is released into one of the azimuthal quadrants. For example, if the entire feed stream is discharged into the reaction medium through two feed holes that are 89 degrees azimuthally separated from each other, in order to determine the azimuthal distribution of 100% of the mass in four azimuthal quadrants. the feed stream is released into the reaction medium in one azimuthal quadrant, since the azimuthal quadrants can be azimuthally oriented so that both feed holes are located in the same azimuthal quadrant.

In addition to the advantages associated with the corresponding azimuthal arrangement of the raw material openings, it has also been found that the corresponding radial arrangement of the raw material openings in the reactor as a bubble column can be of importance. Preferably, a substantial portion of the oxidizable compound introduced into the reaction medium is discharged through the feed openings, which are separated by radial gaps inward from the side wall of the vessel. Therefore, in one embodiment of the present invention, a significant portion of the oxidizable compound enters the reaction zone through feed holes located in the “preferred radial feed zone”, which is separated by a gap inward from the vertical side walls defining the reaction zone.

As for FIG. 7, the preferred radial zone of the feedstock “FZ” may be in the form of a theoretical vertical cylinder centered in the reaction zone 28 and having an outer diameter “D o ” of 0.9 D, where “D” is the diameter of the reaction zone 28. Therefore, the outer the OA annular space, having a thickness of 0.05 D, defines the boundary between the preferred radial zone of the feedstock FZ and the inner side of the side wall defining the reaction zone 28. It is preferable that a small amount of oxidizable compound iv ilos into reaction zone 28 via feed openings to disposed in this outer annulus OA or not introduced at all.

In another embodiment, it is preferred that a small amount of the oxidizable compound is introduced into the center of the reaction zone 28 or not introduced at all. Therefore, as shown in FIG. 8, a preferred radial zone FZ may be in the form of a theoretical vertical annular space centered in the reaction zone, which has an outer diameter D O of 0.9D and an inner diameter of D 1 of 0.2D. Thus, in this embodiment of the invention, the inner cylinder IC having a diameter of 0.2D is “cut out” from the center of the preferred radial zone of the feedstock FZ. Preferably, a small amount of the oxidizable compound is introduced into the reaction zone 28 through the feed openings located in such an inner cylinder IC, or not at all.

In a preferred embodiment of the present invention, a significant portion of the oxidizable compound is introduced into the reaction medium 36 through feed openings located in the preferred radial zone of the feed, regardless of whether the preferred radial feed zone has the cylindrical or ring shape described above. More preferably, at least about 25% of the mass. oxidizable compounds are discharged into the reaction medium 36 through feed openings located in a preferred radial zone of the feed. Even more preferably, at least about 50% of the mass. oxidizable compounds are discharged into the reaction medium 36 through feed openings located in a preferred radial zone of the feed. Most preferably, at least about 75% of the mass. oxidizable compounds are discharged into the reaction medium 36 through feed openings located in a preferred radial zone of the feed.

Although the theoretical azimuthal quadrants and the theoretical preferred feed zone shown in FIG. 7 and 8 are described with reference to the distribution of the liquid-phase feed stream, it is also found that the corresponding azimuthal and radial distribution of the gas-phase flow of the oxidizing agent can also create some advantages. Therefore, in one embodiment of the present invention, the description of the azimuthal and radial distribution of the liquid-phase feed stream presented above is also applicable to the model in which a gas-phase oxidizer stream is introduced into the reaction zone 36.

In FIG. 12 and 13 show an alternative bubble column reactor 200 having a reactor-to-reactor configuration. The bubble column reactor 200 includes an external reactor 202 and an internal reactor 204, wherein the internal reactor 204 is at least partially located in the external reactor 202. In a preferred embodiment, both the external reactor and the internal reactor 202 and 204 are reactors of type of bubble column. Preferably, the external reactor 202 includes an external reaction vessel 206 and an external oxidizer bubbler 208, while the internal reactor 204 includes an internal reaction vessel 210 and an internal oxidizer bubbler 212.

Although FIG. 12 and 13 show the inner reaction vessel 210 as being completely located in the outer reaction vessel 206, it is permissible that the inner reaction vessel 210 is only partially located in the outer reaction vessel 206. However, it is preferable that at least about 50, 90, 95 or 100% of the height of the inner reaction vessel 210 was located in the outer reaction vessel 206. Moreover, it is preferable that a portion of each reaction vessel be raised above a portion of the other reaction vessel, at least about significantly by 0.01, 0.2, 1 or 2 of the maximum diameter of the external reaction vessel.

In a preferred embodiment of the present invention, each of the outer and inner reaction vessels 206 and 210 includes a corresponding vertically arranged side wall, typically having a cylinder configuration. Preferably, the vertically arranged side walls of the outer and inner reaction vessels 206 and 210 are substantially concentric and define an annular space between them. The inner reaction vessel 210 is supported vertically to the outer reaction vessel 206, preferably mainly by means of vertical supports between the lower parts of the respective vessels. In addition, the inner reaction vessel 210 may be supported by the external reaction vessel 206 by a plurality of longitudinal support elements 214 extending between the vertically positioned side wall of the external and internal reaction vessels 206 and 210. Preferably, such longitudinal support elements 214 have a non-contaminated configuration with a minimum flat upside down surface as previously described.

Although it is preferred that the vertically positioned side wall of the inner reaction vessel 210 is substantially cylindrical, it is permissible that some portions of the vertically located side wall of the inner reaction vessel 210 form a concave surface with respect to the adjacent portion of the second reaction zone 218. Preferably, any portion of the vertically located side wall of the inner reaction vessel 218. capacity 210, which is concave relative to the adjacent part of the second reaction zone 218, was less than m approximately 25, 10, 5, or 0.1% of the total surface area of the vertically positioned side wall of the inner reaction vessel 210. Preferably, the ratio of the maximum height of the vertical side wall of the inner reaction vessel 210 to the maximum height of the vertical side wall of the outer reaction vessel 206 is in the range of approximately from 0.1: 1 to 0.9: 1, more preferably in the range of about 0.2: 1 to 0.8: 1, and most preferably in the range of about 0.3: 1 to 0.7: 1.

The external reaction vessel 206 defines here the first reaction zone 216, while the internal reaction vessel 210 defines the second reaction zone 218 here. Preferably, the external and internal reaction vessels 206 and 210 are vertically aligned so that the volumetric center of gravity of the second reaction zone 218 is horizontally offset from the volume center of gravity of the first reaction zone 216, at least about 0.4, 0.2, 0.1 or 0.01 parts of the maximum horizontal diameter of the first reaction zone 216. Preferably about the ratio of the areas of the maximum horizontal cross section of the first reaction zone 216 and the second reaction zone 218 is in the range from about 0.01: 1 to 0.75: 1, more preferably in the range from about 0.03: 1 to 0.5: 1 and most preferably in the range from 0.05: 1 to 0.3: 1. Preferably, the ratio of the horizontal cross-sectional area of the second reaction zone 218 to the horizontal cross-sectional area of the annular space defined between the outer and inner reaction vessels 206 and 210 is at least about 0.2: 1, more preferably in the range of about 0 , 05: 1 to 2: 1, and most preferably in the range of about 0.1: 1 to 1: 1, where the cross-sectional area is measured at ¼ of the height, ½ of the height and / or ¾ of the height of the second reaction zone 218. Preferred at least about 50, 70, 90, or 100% of the volume of the second reaction zone 218 is located in the outer reaction vessel 206. Preferably, the ratio of the volume of the first reaction zone 216 to the volume of the second reaction zone 218 is in the range from about 1: 1 to 100 : 1, more preferably in the range of from about 4: 1 to 50: 1, and most preferably in the range of from 8: 1 to 30: 1. Preferably, the first reaction zone 216 has a ratio of maximum vertical height to maximum horizontal diameter in the range of about 3: 1 to 30: 1, more preferably about 6: 1 to 20: 1, and most preferably in the range of 9: 1 to 15: 1 . Preferably, the second reaction zone 218 has a ratio of maximum vertical height to maximum horizontal diameter in the range of about 0.3: 1 to 100: 1, more preferably in the range of about 1: 1 to 50: 1, and most preferably in the range of 3: 1 up to 30: 1. Preferably, the maximum horizontal diameter of the second reaction zone 218 is in the range of about 0.1 to 5 meters, more preferably in the range of about 0.3 to 4 m, and most preferably in the range of 1 to 3 m. Preferably, the maximum vertical height of the second reaction zone is 218 is in the range of about 1 to 100 m, more preferably in the range of about 3 to 50 m, and most preferably in the range of 10 to 30 m. The ratio of maximum horizontal d the meter of the second reaction zone 218 to the maximum horizontal diameter of the first reaction zone 216 is in the range of from about 0.05: 1 to 0.8: 1, more preferably in the range of from about 0.1: 1 to 0.6: 1, and most preferably in the range from 0.2: 1 to 0.5: 1. Preferably, the ratio of the maximum vertical height of the second reaction zone 218 to the maximum vertical height of the first reaction zone 216 is in the range of about 0.03: 1 to 1: 1, more preferably in the range of about 0.1: 1 to 0.9: 1, and most preferably in the range of 0.3: 1 to 0.8: 1. Any parameters (e.g., height, width, area, volume, relative horizontal placement and relative vertical placement) defined here for the external reaction vessel 206 and the components are also considered applicable to the first reaction zone 216 defined by the external reaction vessel 206, and vice versa . In addition, any parameters defined here for the inner reaction vessel 210 and the constituent parts are also considered applicable to the second reaction zone 218 defined by the internal reaction vessel 210, and vice versa.

When the reactor is operated as a bubble column 200, the multiphase reaction medium 220 is first oxidized in the first reaction zone 216 and then oxidized in the second reaction zone 218. Therefore, during normal operation, the first part of the reaction medium 220a is located in the first reaction zone 216, then as the second part of the reaction medium 220b is located in the second reaction zone 218. After processing in the second reaction zone 218, the suspension phase (i.e., liquid and solid phases) of the reaction medium 220b is withdrawn from Torah reaction zone 218 and discharged from the reactor type bubble column 200 through an outlet 222 for the slurry for further processing downstream side.

The internal reactor 204 preferably has at least one internal gas hole that allows additional molecular oxygen to be introduced into the second reaction zone 218. Preferably, the plurality of internal gas holes are defined by the internal oxidizer bubbler 212. The descriptions for the oxidizer bubbler 34 in FIG. 1-5 are also applicable to the internal oxidizer bubbler 212 in terms of piping dimensions and configurations, orifice size and configuration, working differential pressures, and liquid flushing. A noticeable difference is that it is preferable to position the oxidizer bubbler 212 relatively higher in order to use the lower part of the inner reaction vessel 210 as a deaeration zone. For example, the embodiments described in the present invention for the oxidation of para-xylene to form TFA (TPA) provide significantly reduced spatio-temporal reaction rates near the bottom of the second reaction zone 218, and this reduces the effect of deaeration on the formation of impurities. The inner reaction vessel 210 has a maximum height of "H i ". Preferably, at least about 50, 75, 95, or 100% of the total cross-sectional area defined by all internal gas openings is at least 0.05H i , 0.1H i, or 0.25H i from the top of the inner reaction vessel 210. It is also preferred that at least about 50, 75, 95, or 100% of the total cross-sectional area defined by all the internal gas openings is at least about 0.5H i , 0.25H i or 0.1H i above the bottom of the inner reaction vessel 21 0. Preferably, at least about 50, 75, 95, or 100% of the total cross-sectional area defined by all internal gas openings is at least 1, 5, or 10 m from the top of the internal reaction vessel 210 and at a distance of at least about 0.5, 1 or 2 m from the bottom of the inner reaction vessel 210. It is preferable that at least about 50, 75, 95 or 100% of the total cross-sectional area defined by all internal gas holes, connected directly but with the second reaction zone 218, it was not directly connected to the first reaction zone 216. As used in this case, the definition of “cross-sectional area” means the minimum surface area (flat or non-planar) that will cover the hole.

In the General case, the model by which the flows of raw materials, oxidizing agent and reflux are introduced into the external reactor 202, and the model by which the external reactor 202 operates, are essentially the same as described above for the bubble column reactor 20 in FIG. 1-11. However, one difference between the external reactor 202 (FIGS. 12 and 13) and the bubble column reactor 20 (FIGS. 1-11) is that the external reactor 202 does not include an outlet that allows the suspension phase of the reaction medium 220a to directly discharge from an external reaction vessel 206 for downstream processing. More specifically, a bubble column reactor 200 requires that the suspension phase of the reaction medium 220a, before being discharged from the reactor as a bubble column 200, first passes through the internal reactor 204. As mentioned above, in the second reaction zone 218 of the internal reactor 204, the reaction medium 220b is further oxidized to facilitate purification of the liquid and / or solid phases of the reaction medium 220b.

In a method where para-xylene is introduced into reaction zone 216, the liquid phase of reaction medium 220a, which leaves the first reaction zone 216 and enters the second reaction zone 218, typically contains at least some para-toluic acid. Preferably, a significant portion of the para-toluic acid entering the second reaction zone is oxidized in the second reaction zone 218. Therefore, it is preferable that the time-averaged concentration of para-toluic acid in the liquid phase of the reaction medium 220b leaving the second reaction zone 218 was less than the time-averaged concentration of para-toluic acid in the liquid phase of the reaction medium 220a / b entering the second reaction zone 218. It is preferable that the time-averaged concentration of pa p-toluic acid in the liquid phase of the reaction medium 220b leaving the second reaction zone 218 was less than about 50, 10 or 5% of the time-averaged concentration of p-toluic acid in the liquid phase of the reaction medium 220a / b entering the second reaction zone 218. Preferably, the time-averaged concentration of para-toluic acid in the liquid phase of the reaction medium 220a / b entering the second reaction zone 218 is at least about 250 mass. hours / million, more preferably is in the range from about 500 to 6000 mass. hours / million and most preferably in the range from 1000 to 4000 mass. ppm Preferably, the time-averaged concentration of para-toluic acid in the liquid phase of the reaction medium 220b leaving the second reaction zone 218 is less than about 1000, 250 or 50 mass. ppm

The inner reaction vessel 210 is equipped with at least one straight hole that allows the reaction medium 220a / b to pass directly between the reaction zone 216 and the second reaction zone 218. It is preferred that substantially all straight openings in the inner reaction vessel 210 are located near the top internal reaction vessel 210. Preferably, at least about 50, 75, 90, or 100% of the total cross-sectional area defined by all straight openings is less than four i about 0.5H, 0.25H or 0.1H i i from the top of the inner reaction container 210. Preferably, less than about 50, 25, 10, or 1% of the total opening area defined by all straight holes in the inner reaction vessel 210 is more than about 0.5 N i , 0.25 N i or 0.1 N i from the top of the inner reaction vessel 210. Most preferably, the straight hole defined by the inner reaction vessel 210 is the only upper hole 224 located at the uppermost end of the inner a reaction vessel 210. The ratio of the cross-sectional area of the upper opening 224 to the maximum horizontal cross-sectional area of the second reaction zone 218 is preferably at least 0.1: 1, 0.2: 1 or 0.5: 1.

During normal operation of the reactor as a bubble column 200, the reaction medium 220 passes from the first reaction zone 216, through the direct hole (s) (for example, the upper hole 224) into the inner reaction vessel 210 and into the second reaction zone 218. In the second reaction zone zone 218, the suspension phase of the reaction medium 220b is usually moved downward through the second reaction zone 218, while the gas phase of the reaction medium 220b is usually moved upward. Preferably, the inner reaction vessel 210 defines at least one discharge opening that allows the suspension phase to exit the second reaction zone 218. The suspension phase leaving the discharge opening of the internal reaction vessel 210 then exits the reactor as a bubble column 200 through an outlet for slurry 222. Preferably, the discharge opening is located on or near the bottom of the inner reaction vessel 210. Preferably, at least about 50, 75, 90 or 100% of the total area of the cross section crossing, certain of paged by all holes in the inner reaction vessel 210 is disposed within about 0.5H i, 0.25H or 0.1H i i, from bottoms of the inner reaction vessel 210.

Since the reaction medium 220b is processed in the second reaction zone 218 of the internal reactor 204, it is preferable to lower the gas retention of the reaction medium 220b while the suspension phase of the reaction medium 220b flows downward through the second reaction zone 218. Preferably, the time-averaged gas retention ratio of the reaction medium 220a / b entering the second reaction zone 218 to the gas retention of the reaction medium 220b leaving the second reaction zone 218 is at least about 2: 1, 10: 1 or 25: 1. Preferably, the time-averaged gas retention of the reaction medium 220a / b exiting the second reaction zone 218 is in the range of about 0.4 to 0.9, more preferably in the range of about 0.5 to 0.8, and most preferably in the range from 0.55 to 0.7. Preferably, the time-averaged gas retention of the reaction medium 220b leaving the second reaction zone 218 is less than about 0.1, 0.05, or 0.02. Preferably, the ratio of the time-averaged gas retention of the reaction medium 220a in the first reaction zone 216 to the gas retention of the reaction medium 220b in the second reaction zone 218 is greater than about 1: 1, more preferably in the range of about 1.25: 1 to 5: 1 and most preferably in the range from 1.5: 1 to 4: 1, where the gas retention value is measured at any height of the first and second reaction zones 216 and 218, at any corresponding heights of the first and second reaction zones 216 and 218, ¼ of the height of the first and / or the second reaction zones 216 and 218, ½ the height of the first and / or second reaction zones 216 and 218, ¾ the height of the first and / or second reaction zones 216 and 218, and / or represents the average value over all the heights of the first and / or second reaction zones 216 and 218. Preferably, the time-averaged gas retention of part of the reaction medium 220a in the first reaction zone 216 is in the range of about 0.4 to 0.9, more preferably in the range of about 0.5 to 0.8, and most preferably in the range from 0.55 to 0.70, where gas retention is measured at any height of the first reaction zone 216, ¼ the height of the first reaction zone 216, ½ the height of the first reaction zone 216, ¾ the height of the first reaction zone 216, and / or is the average over the entire height of the first reaction zone 216. Preferably, the time-averaged retention the gas portion of the reaction medium 220b in the second reaction zone 218 is in the range of from about 0.01 to 0.6, more preferably in the range of from about 0.03 to 0.3, and most preferably in the range of from 0.08 to 0.2, where gas retention is measured at any the first height of the second reaction zone 218, ¼ the height of the second reaction zone 218, ½ the height of the second reaction zone 218, ¾ the height of the second reaction zone 218, and / or is the average over the entire height of the second reaction zone 218.

The temperature of the reaction medium 220 is preferably approximately equal to the temperature in the first and second reaction zones 216 and 218. Preferably, this temperature is in the range of about 125 to 200 ° C, more preferably in the range of about 140 to 180 ° C, and most preferably in the range of 150 to 170 ° C. However, temperature differences occur within the first reaction zone 216, which is the same as that described with reference to FIG. 28. Preferably, temperature differences of the same order also exist within the second reaction zone 218, as well as between the first reaction zone 216 and the second reaction zone 218. Such additional temperature gradients are associated with the chemical reaction taking place in the second reaction zone 218, with the introduction of an additional oxidizing agent into the second reaction zone 218 and the static pressures existing in the second reaction zone 218, compared with pressures in the first reaction zone 216. As described above, the retention of bubbles preferably higher in the first reaction zone 216 than in the second reaction zone 218. Therefore, at elevations below the upper opening 224, the static pressure in the reaction zone 216 is greater than in the second reaction zone 218. The magnitude of this pressure difference depends on the density of the liquid or suspension and from the difference in the retention of bubbles between the two reaction zones. The magnitude of this pressure difference increases at elevations even lower than the upper opening 224.

In one embodiment of the present invention, part of the oxidizable compound (e.g., para-xylene) fed to the reactor as a bubble column 200 is introduced directly into the second reaction zone 218 of the internal reactor 204. However, at least about 90 95, 99 or 100 mol% the total oxidizable compound fed to the reactor as a bubble column 200 is introduced into the first reaction zone 216 (and not into the second reaction zone 218). Preferably, the molar ratio of the amount of oxidizable compound introduced into the first reaction zone 216 to the amount of oxidizable compound introduced into the second reaction zone 218 is at least 2: 1, 4: 1 or 8: 1.

Although FIG. 12 and 13 show the configuration when a portion of the total molecular oxygen supplied to the bubble column reactor 200 is introduced into the second reaction zone 218 of the inner reactor 204 through the inner oxidizer bubbler 212, it is preferred that the bulk of the total molecular oxygen fed to the reactor is of the type the bubble column 200 was introduced into the first reaction zone 216, with the remaining amount being introduced into the second reaction zone 218. Preferably, at least about 70, 90, 95 or 98 mol%. of the total molecular oxygen fed to the reactor as a bubble column 200 is introduced into the first reaction zone 216. Preferably, the molar ratio of the amount of molecular oxygen introduced into the first reaction zone 216 to the amount of molecular oxygen introduced into the second reaction zone 218 is at least at least about 2: 1, more preferably in the range of about 4: 1 to 200: 1, most preferably in the range of 10: 1 to 100: 1. Although it is possible that a certain amount of solvent and / or an oxidizable compound (e.g., para-xylene) is fed directly to the second reaction zone 218, it is preferred that less than about 10, 5, or 1% of the mass. the total amount of solvent and / or oxidizable compound supplied to the reactor as a bubble column 200 was fed directly to the second reaction zone 218.

The volume, residence time and space velocity of the medium 220a in the first reaction zone 216 of the external reaction vessel 206 are preferably significantly higher than the volume, residence time and space velocity of the reaction medium 220b in the second reaction zone 218 of the internal reaction vessel 210. Therefore, the bulk of the oxidizable compound ( for example, para-xylene) fed to the reactor as a bubble column 200 is preferably oxidized in the first reaction zone 216. Preferably, at least about 80, 90 or 95% of the mass. of the total oxidizable compound, which is oxidized in a bubble column reactor 200, is oxidized in the first reaction zone 216. Preferably, the time-averaged reduced gas velocity of the reaction medium 220a in the first reaction zone 216 is at least about 0.2 , 0.4, 0.8 or 1 meter per second, where the reduced gas velocity is measured at any height of the first reaction zone 216, ¼ the height of the first reaction zone 216, ½ the height of the first reaction zone 216, ¾ the height of the first reaction zone 216 , and / or n edstavlyaet an average over the entire height of the first reaction zone 216.

Although the reaction medium 220b in the second reaction zone 218 may have the same reduced gas velocity as the reaction medium 220a in the first reaction zone 216, it is preferable that the time-averaged reduced gas velocity of the reaction medium 220b in the second reaction zone 218 is less than the time-averaged and a volume-averaged reduced gas velocity of the reaction medium 220b in the second reaction zone 218. This reduced reduced gas velocity in the second reaction zone 218 is made possible, for example, by reducing molecular oxygen consumption in the second reaction zone 218 compared to the first reaction zone 216. Preferably, the ratio of the time-averaged reduced velocity of the reaction medium 220a in the first reaction zone 216 to the time-averaged reduced velocity of the reaction medium 220b in the second reaction zone 218 is at least , approximately 1.25: 1, 2: 1 or 5: 1, where the reduced gas velocities are measured at any height of the first and second reaction zones 216 and 218, at any corresponding heights of the first and second reaction zones 216 and 218, by ¼ the height of the first and / or second reaction zones 216 and 218, by ½ the heights of the first and / or second reaction zones 216 and 218, by ¾ the heights of the first and / or second reaction zones 216 and 218, and / or represents is the average value over all heights of the first and / or second reaction zones 216 and 218. Preferably, the time-averaged and volume-averaged reduced gas velocity of the reaction medium 220b in the second reaction zone 218 is less than about 0.2, 0.1, or 0, 06 meters per second, where the reduced gas velocity is measured at any height by the second reaction zone 218, ¼ the height of the second reaction zone 218, ½ the height of the second reaction zone 218, ¾ the height of the second reaction zone 218, and / or is the average over the entire height of the second reaction zone 218. With such low reduced gas velocities, a downward flow of the suspension phase of the reaction medium 220b in the second reaction zone 218 is obtained in order to switch directionally to the piston mode of the two-phase flow. For example, during the oxidation of para-xylene to TPA, the relative vertical concentration gradients in the liquid phase of para-toluic acid can be much larger in the second reaction zone 218 than in the first reaction zone 216. And this despite the fact that the second reaction zone 218 represents a bubble column having an axial displacement of the liquid and suspension compositions. The time-averaged reduced velocity of the suspension phase (solid + liquid) and the liquid phase of the reaction medium 220b in the second reaction zone 218 is preferably less than about 0.2, 0.1, or 0.06 m / s, where the reduced velocity is measured at any the height of the second reaction zone 218, ¼ the height of the second reaction zone 218, ½ the height of the second reaction zone 218, ¾ the height of the second reaction zone 218, and / or is the average over the entire height of the second reaction zone 218.

In one embodiment of the present invention, the bubble column reactor 200 is operated so as to provide solids deposition in the internal reactor 204. If solids deposition is desired, it is preferred that the time-averaged and volume-averaged reduced gas velocity of the reaction medium 220b during the second reaction zone was less than about 0.05, 0.03, or 0.01 m / s. Furthermore, if solids deposition is desired, it is preferable that the time-averaged and volume-averaged reduced velocity of the suspension phase and liquid phase of the reaction medium 220b in the second reaction zone 218 is less than about 0.01, 0.005, or 0.001 m / s.

Although it is permissible that some portion of the slurry phase exiting the internal reactor 204 be directly recycled back to the first reaction zone 216 without additional downstream processing, it is preferred that the reaction medium 220b is directly recycled from the lower elevation marks of the second reaction zone 218 to the first reaction zone 216 has been minimized. Preferably, the mass of the reaction medium 220b (solid, liquid, and gas phases) leaving the lower 25% of the volume of the second reaction zone 218 and directly recycled back to the first reaction zone 216 without further downstream processing is less than 10, 1, or 0.1 parts from the mass (solid, liquid and gas phases) of the reaction medium 220b leaving the second reaction zone 218 and then subjected to a downward flow treatment. Preferably, the mass of the reaction medium 220b leaving the lower 50% of the volume of the second reaction zone 218 and directly recycled back to the first reaction zone 216 without additional downstream processing is less than 20, 2 or 0.2 parts of the mass of the reaction medium 220b leaving from the second reaction zone 218 and then processed in a downward flow. Preferably less than about 50, 75 or 90% of the mass. the liquid phase of the reaction medium 220b leaving the second reaction zone 218 through the openings in the lower 90, 60, 50 or 5% of the volume of the second reaction zone 218, is introduced into the first reaction zone 216 within 60, 20, 5 or 1 min after leaving the second reaction zone 218. Preferably, the liquid phase of the reaction medium 220b located in the second reaction zone 218 has a mass-average residence time in the second reaction zone 218 of at least about 1 minute, more preferably is in the range of about 2 to 60 minutes, and most pre respectfully in the range of 5 to 30 minutes Preferably less than about 50, 75 or 90% of the mass. the liquid phase of the reaction medium 220a / b introduced into the second reaction zone 218 is included in the second reaction zone 218 in the lower 90, 60 or 30% of the volume of the second reaction zone 218. Preferably, less than about 50, 75 or 90% of the mass. of the total volume of the liquid phase of the reaction medium 220a / b introduced as a liquid phase stream into the first reaction zone 216, enters the first reaction zone 216 within 60, 20, 5, or 1 min after removal from the second reaction zone 218 through the outlet for suspension 222. Preferably, at least about 75, 90, 95 or 99% of the mass. the liquid phase of the reaction medium 220b withdrawn from the second reaction zone 218 leaves the second reaction zone 218 through openings in the lower 90, 60, 30, or 5% of the volume of the second reaction zone 218.

A bubble column reactor design in a reactor-in-reactor configuration 200 can be modified in various ways without departing from the scope of the present invention. For example, the internal reaction vessel 210 may have a greater height than the external reaction vessel 206 if the internal reaction vessel 210 extends below the lower end of the external reaction vessel 206. The external and internal reaction vessels 206 and 210 may be cylindrical, as shown, or may have any form. The outer and inner reaction vessels 206 and 210 need not be axisymmetric, axially vertical, or concentric. The gas phase exiting the internal reactor 204 can be directed outward from the reactor as a bubble column 200 without being combined with the reaction medium 220a in the first reaction zone 216. However, from the point of view of ignition safety, it is desirable to limit the volumes of trapped gas pockets to less than about 10 , 2 or 1 m 3 . In addition, it is not necessary that the slurry phase leaving the internal reactor 204 exit through a single slurry hole in the bottom of the inner reaction vessel 210. The slurry phase may exit the reactor as a bubble column 200 through a lateral outlet in a pressurized side wall external reactor 202.

In FIG. 14 shows a bubble column reactor 300 having a reactor-in-reactor configuration and a step diameter configuration. The bubble column reactor 300 includes an external reactor 302 and an internal reactor 304. The external reactor 302 is an external reaction vessel 306 having a wide lower section 306a and a narrow upper section 306b. Preferably, the diameter of the narrow upper section 306b is smaller than the diameter of the wide lower section 306a. With the exception of the configuration of the stepped diameter of the external reaction vessel, the reactor is a bubble column 300 of FIG. 14 is preferably configured and operates substantially similar to a bubble column reactor 200 of FIG. 12 and 13 described above.

In FIG. 15 shows a reactor system 400 including a primary oxidation reactor 402 and a secondary oxidation reactor 404. The primary oxidation reactor 402 is preferably configured and operates substantially similar to the external reactor 202 of FIG. 12 and 13. The secondary oxidation reactor 404 is preferably configured and operates substantially similarly to the internal reactor 204 of FIG. 12 and 13. However, the main difference between the reactor system 400 of FIG. 15 and a bubble column reactor 200 of FIG. 12 and 13, the secondary oxidation reactor 404 of the reactor system 400 is located outside the primary oxidation reactor 402. In the reactor system 400 of FIG. 15, an inlet conduit 405 is used to transfer part of the reaction medium 420 from a primary oxidation reactor 402 to a secondary oxidation reactor 404. In addition, an exhaust conduit 407 is used to transfer head gases from the top of the secondary oxidation reactor 404 to the primary oxidation reactor 402.

During normal operation of the reaction system 400, the reaction medium 420 is first oxidized in the primary reaction zone 416 of the primary oxidation reactor 402. The reaction medium 420a is then withdrawn from the primary reaction zone 416 and transferred to the secondary reaction zone 418 via line 405. In the secondary reaction zone 418, liquid and / or the solid phases of the reaction medium 420b are further oxidized. Preferably, at least about 50, 75, 95 or 99% of the mass. the liquid phase and / or the solid phase was discharged from the primary reaction zone 416 for processing in the secondary reaction zone 416. The head gases leave the upper gas outlet of the secondary oxidation reactor 404 and are transferred back to the primary oxidation reactor 402 via line 407. The suspension phase of the reaction medium 420b exits the lower outlet of the slurry 422 of the secondary oxidation reactor 404 and then undergoes additional downstream processing.

The inlet conduit 405 may be connected to the primary oxidation reactor 402 at any height. Although in FIG. 15 is not shown, the reaction medium 420 may be mechanically pumped into the secondary reaction zone 418, if desired. However, it is more preferable to use hydrostatic pressure (gravity) to transfer the reaction medium 420 from the primary reaction zone 416 through the inlet pipe 405 and to the secondary reaction zone 418. Accordingly, it is preferable that the inlet pipe 405 is connected at one end to the upper 50, 30, 20 or 10% of the total height and / or total volume of the primary reaction zone 416. Preferably, the other end of the inlet conduit 405 is connected to the upper 30, 20, 10, or 5% of the total height and / or total volume of the secondary reaction zone 418. Preferably, the inlet conduit 405 is horizontal and / or inclined downward from the primary oxidation reactor 402 towards the secondary oxidation reactor 404. The outlet conduit 407 may be connected to any elevation in the secondary oxidation reactor 404, but it is preferred that the outlet conduit 407 is connected to the secondary reactor oxidation 404 is above the attachment height of the intake pipe 405. More preferably, the exhaust pipe 407 is attached to the top of the secondary oxidation reactor 404. The outlet conduit 407 is preferably connected to the primary oxidation reactor 402 above the attachment height of the inlet conduit 405. More preferably, the outlet conduit 407 is attached to the upper 30, 20, 10, or 5% of the total height and / or total volume of the primary reaction zone 416. Preferably, the outlet conduit 407 is horizontal and / or inclined downward from the secondary oxidation reactor 404 in the direction of the primary oxidation reactor 402. Although in FIG. 15 is not shown, the exhaust pipe 407 can also be attached directly to a gas exhaust pipe that discharges a gaseous effluent from the top of the primary oxidation reactor 402. The upper space of the secondary reaction zone 416 may be above or below the upper space of the primary reaction zone 418. More preferably the upper length of the primary reaction zone 416 is in the range from 10 m above and up to 50 m below, from 2 to 40 m below, from 5 to 30 m below the upper length of the secondary reaction zone 418. N the lower outlet for slurry 422 may extend from any elevation of the secondary oxidation reactor 404, but it is preferable that the lower outlet for slurry 422 is connected to the secondary oxidation 404 below the elevation of the inlet 405. The attachment point of the lower outlet for slurry 422 more preferably with a large distance is separated in height from the connection point of the intake pipe 405, and the two connections are separated by at least an approximate Only 50, 70, 90, or 95% of the height of the secondary reaction zone 418. More preferably, the bottom outlet for slurry 422 is attached to the bottom of the secondary oxidation reactor 404, as shown in FIG. 15. The lower length of the secondary reaction zone 418 can be set at a height above or below the lower length of the primary reaction zone 416. More preferably, the lower length of the primary reaction zone 416 is at an altitude of 40, 20, 5 or 2 m above or below the lower length of the secondary reaction zone 418.

The parameters (e.g., height, width, area, volume, relative horizontal placement and relative vertical placement) defined here for the primary oxidation reactor 402 and the components are also considered applicable to the primary reaction zone 416 defined by the primary oxidation reactor 402, and vice versa. Any parameters defined in the present invention for the secondary oxidation reactor 404 and its constituents are also considered applicable to the secondary reaction zone 418 defined by the secondary oxidation reactor 404, and vice versa.

As mentioned above, it is preferable that the secondary oxidation reactor 404 is located outside the primary oxidation reactor 402. Preferably, the secondary oxidation reactor 404 is located along the side of the primary oxidation reactor 402 (i.e., at least a portion of the primary and secondary oxidation reactors 402 and 404 share a common elevation mark). The primary reaction zone 416 of the primary oxidation reactor 402 has a maximum diameter of "D p ". The volumetric center of gravity of the secondary reaction zone 418 is preferably horizontally separated from the volumetric center of gravity of the primary reaction zone 416 by at least about 0.5D p , 0.75D p or 1.0D p and less than about 30D p , 10D p or 3D p .

In FIG. 16, a reactor system 500 is shown including a primary oxidation reactor 502 and a secondary oxidation reactor 504. In this case, the primary oxidation reactor defines a primary oxidation zone 516, while the secondary oxidation reactor 504 defines a secondary oxidation zone 518. Each reaction zone 516 and 518 receives a part of the reaction Wednesday 520.

The configuration and operation of the reactor system 500 (FIG. 16) is preferably similar to the configuration and operation of the reactor system 400 (FIG. 15). However, in the reactor system 500, the vertically mounted side wall of the primary oxidation reactor 502 defines the boundaries of at least one enlarged opening 505, which allows the reaction medium 520 to be transferred from the primary reaction zone 516 to the secondary reaction zone 518, while simultaneously allowing the separated gas to be transferred phase from the secondary reaction zone 518 to the primary reaction zone 516. Preferably, the cross-sectional area of the enlarged hole 505 divided by the area of maximum g rizontalnogo vertical cross-sectional portion of the secondary reaction zone 518, is in the range from about 0.01 to 2, 0.02 to 0.5 or from 0.04 to 0.2. The primary reaction zone 516 of the primary oxidation reactor 502 has a maximum height of "N p ". Preferably, the center of the enlarged opening 505 is vertically separated by a gap of at least about 0.1 N p , 0.2 N p or 0.3 N p from the top and / or bottom of the primary reaction zone 516.

In FIG. 17-25 show a series of bubble column reactors equipped with internal structures having various configurations. It has been found that the use of one or more internal structures surrounded by a reaction medium unexpectedly modifies through mixing of the reaction medium. The internal structure defines a static zone having reduced turbulence compared with the turbulence of the reaction medium surrounding the static zone.

As shown in FIG. 17-25, the internal structure may take various forms. In particular, FIG. 17 illustrates a bubble column reactor 600 in which a generally cylindrical internal structure 602 is used to define a static zone. The internal structure 602 is essentially centered in the main reaction zone of the reactor as a bubble column 600 and is separated by a vertical gap from the upper and lower ends of the main reaction zone. FIG. 18 illustrates a bubble column reactor 610 that utilizes a generally cylindrical internal structure 612 similar to the internal structure 602 of FIG. 17. However, the internal structure 612 in FIG. 18 is not centered in the main reaction zone of the reactor as a bubble column 610. Moreover, the volume center of gravity of the static zone defined by the internal structure 612 is horizontally offset from the volume center of gravity of the main reaction zone. In addition, the bottom of the internal structure 612 is located near the lower tangent line of the reactor as a bubble column 610. FIG. 19 illustrates a bubble column reactor 620 that utilizes a generally cylindrical internal structure 622 that is higher than the internal structure 602 and 612 of FIG. 17 and 18. In addition, the volumetric center of gravity of the static zone of the internal structure 622 is offset from the volumetric center of the main reaction zone of the reactor by the type of bubble column 620. FIG. 20 illustrates a bubble column reactor 630 that utilizes an internal structure including a generally cylindrical upper portion 632 and typically a cylindrical lower portion 634. The lower portion 634 of the internal structure has a narrower diameter than the upper portion 632. FIG. 21 illustrates a bubble column reactor 640 using an internal structure including a generally cylindrical lower portion 642 and typically a cylindrical upper portion 644. The upper portion 644 of the internal structure has a narrower diameter than the lower portion 642. FIG. 22 illustrates a bubble column reactor 650 using the first, second, and third separate internal structures 652, 654, and 656. The internal structures 652, 654, and 656 are separated by vertical gaps from each other. The volume centers of gravity of the static zones defined by the first and third internal structures 652 and 656 are horizontally aligned with the volume center of gravity of the main reaction zone of the reactor as a bubble column 650. However, the volume center of gravity of the static zone defined by the internal structure 654 is horizontally offset from the volume center of gravity the main reaction zone of the reactor as a bubble column 650. FIG. 23 illustrates a bubble column reactor 660 in which a pair of side-by-side first and second internal structures 662 and 664 are used. The volume centers of gravity of the static zones defined by the first and second internal structures 662 and 664 are horizontally spaced apart and horizontally separated by a gap from the volume center of gravity of the main reaction zone of the reactor as a bubble column 660. In addition, the first and second internal structures 662 and 664 have a side-by-side configuration such that at least m Here, part of the first and second internal structures 662 and 664 share a common elevation mark. FIG. 24 illustrates a bubble column reactor 670 that typically uses a prismatic internal structure 672. In particular, the internal structure 672 has a generally triangular horizontal cross section. FIG. 25 illustrates a bubble column reactor 680 that typically uses a cylindrical internal structure 682 that is smaller than the internal structure 602 of FIG. 17. However, the external reaction vessel of the reactor as a bubble column 680 has a stepped diameter formed by a narrow lower section 682 and a wide upper section 684.

As shown in FIG. 17-25, the internal structure used in accordance with one embodiment of the present invention may have a number of configurations and may be located at different positions within the main reaction zone of the reactor as a bubble column. In addition, the internal structure and the static zone defined by it can be obtained from many different materials. In one embodiment of the present invention, the internal structure is completely closed so that no environmental reaction medium enters the internal structure. Such a closed internal structure may be hollow or solid. In another embodiment of the present invention, the internal structure includes one or more openings that allow the reaction medium to enter a static zone defined by the internal structure. However, since one of the tasks of the static zone is to create a zone of reduced turbulence with respect to the turbulence of the reaction medium surrounding the internal structure, it is preferable that the internal structure does not allow a significant amount of the reaction medium to quickly flow through the internal structure.

The specific configuration and technological parameters of a bubble column reactor equipped with one or more internal structures will be described in more detail. Preferably, the internal structure is located completely inside the external reaction vessel of the reactor as a bubble column; however, it is permissible that at least a portion of the internal structure protrude beyond the outer reaction vessel of the reactor as a bubble column. As mentioned above, while the reactor is operating as a bubbler, the internal structure defines at least one static zone within the reactor as a bubbler. The main reaction zone of the reactor as a bubble column and the static zone have separate volumes (that is, they do not overlap one another). The main reaction zone of the reactor as a bubble column is defined inside the external reaction vessel of the reactor as a bubble column, but outside the internal structure.

As mentioned above, the static zone defined by the internal structure is a volume that has reduced turbulence with respect to the turbulence of the adjacent reaction medium in the main reaction zone. Preferably, at least about 90, 95, 98, or 99.9% of the volume of the static zone is filled with material other than the reaction medium, and / or is filled with a part of the reaction medium having substantially reduced turbulence compared to the turbulence of the reaction environment adjacent to the internal structure. If the static zone comprises any part of the reaction medium, it is preferable that the part of the reaction medium located in the static zone has a mass-average residence time in the static zone of at least about 2, 8, 30, or 120 minutes. If the static zone includes any part of the reaction medium, it is preferred that the time-averaged gas retention of the reaction medium in the static zone is less than about 0.2, 0.1, 0.5, or 0.01, where gas retention is measured at any elevation static zone, ¼ the height of the static zone, ½ the height of the static zone, ¾ the height of the static zone, and / or represents the average over the entire height of the static zone. Preferably, the time-averaged gas retention of the reaction medium in the reaction zone is in the range of from about 0.2 to 0.9, more preferably from about 0.5 to 0.8, and most preferably from 0.55 to 0.7, where gas retention is measured at any elevation of the reaction zone, ¼ the height of the reaction zone, ½ the height of the reaction zone, ¾ the height of the reaction zone, and / or is the average over the entire height of the reaction zone. If the static zone includes any part of the reaction medium, it is preferable that the time-averaged reduced gas velocity of the reaction medium in the static zone is less than about 0.4, 0.2, 0.1, or 0.05 m / s, where the reduced gas velocity measured at any height of the static zone, ¼ the height of the static zone, ½ the height of the static zone, ¾ the height of the static zone, and / or is the average over the entire height of the static zone. Preferably, the time-averaged reduced gas velocity of the reaction medium in the reaction zone is at least about 0.2, 0.4, 0.8 or 1 m / s, where the reduced gas velocity is measured at any elevation of the reaction zone , ¼ the height of the reaction zone, ½ the height of the reaction zone, ¾ the height of the reaction zone, and / or represents the average over the entire height of the reaction zone. If the static zone includes any part of the reaction medium, it is preferable that the time-averaged reduced velocity of the liquid phase of the reaction medium in the static zone is less than about 0.04, 0.01 or 0.004 m / s, where the reduced velocity of the liquid phase is measured at any point the height of the static zone, by ¼ the height of the static zone, ½ the height of the static zone, на the height of the static zone, and / or represents the average over the entire height of the static zone. Preferably, the time-averaged reduced velocity of the liquid phase of the reaction medium in the reaction zone is less than about 0.1, 0.04 or 0.01 m / s, where the reduced velocity of the liquid phase is measured at any elevation of the reaction zone, by ¼ of the height the reaction zone, by ½ the height of the reaction zone, by ¾ the height of the reaction zone, and / or represents the average over the entire height of the reaction zone. Any parameters (e.g., height, width, area, volume, relative horizontal placement and relative vertical placement) defined here for the internal structure are also considered applicable to the static zone defined by the internal structure, and vice versa.

Preferably, the size of the static zone defined by the internal structure is such that the static zone has at least one location that is spaced apart from the reaction zone by at least about 0.05 part of the maximum horizontal diameter of the reaction zone or approximately 0.2 m, sometimes more. Preferably, the static zone has at least one location that is spaced apart from the reaction zone by at least about 0.4, 0.7 or 1.0 m. Preferably, the static zone has at least one location, which is separated by a gap from the reaction zone, at least about 0.1, 0.2 or 0.3 parts from the maximum horizontal diameter of the reaction zone. The static zone preferably has at least two locations that are separated from each other by a vertical distance that is at least about 0.5, 1.2, or 4 parts of the maximum horizontal diameter of the reaction zone. Preferably, said two vertically spaced apart locations in the static zone are also each separated from the reaction zone by at least about 0.05, 0.1, 0.2 or 0.3 parts of the maximum horizontal diameter of the reaction zone. Preferably, said two vertically spaced apart locations in the static zone are separated from each other by a vertical span of at least about 1, 3, 10, or 20 m, and each is also separated from the reaction zone by at least about 0.1, 04 0.7 or 1 m. Preferably, the volume of the static zone is in the range of about 1 to 50% of the volume of the main reaction zone, more preferably in the range of about 2 to 25% of the volume of the main reaction zone, and most preferably in vomited from 4 to 15% by volume of the primary reaction zone.

The external reaction vessel of the reactor as a bubble column includes, as a rule, a cylindrical vertical external side wall. Preferably, the inner structure includes a cylindrical vertical side wall that is spaced inward inward from the outer side wall. Preferably, the internal structure is not part of the heat exchanger. Therefore, it is preferable that the time-averaged specific heat flux through vertically mounted inner side walls of the internal structure be less than about 100, 15, 3, or 0.3 kilowatts per square meter. The annular space filled with the reaction medium is preferably limited between the inner and outer side walls. The internal structure is supported vertically to the external vessel, preferably with vertically mounted supports between the lower parts of the internal structure and the lower part of the external reaction vessel. Furthermore, the internal structure is preferably supported by an external reaction vessel due to a plurality of non-contaminated longitudinal support elements extending inward from the outer side wall to the inner side wall. Preferably, the horizontal cross-sectional area of the static zone by ¼ height, ½ height and / or ¾ of the height of the static zone is at least about 2, from 5 to 75, or from 10 to 30% of the horizontal cross-sectional area of the annular space at the corresponding height mark . Preferably, the maximum height of the outer vertical side wall is in the range of about 10 to 90% of the maximum height of the inner vertical side wall, more preferably in the range of about 20 to 80% of the maximum height of the outer vertical side wall, and most preferably in the range of about 30 up to 70% of the maximum height of the outer vertical side wall. Although it is preferable that the inner side wall has a generally cylindrical configuration, it is permissible that a part of the inner side wall be concave relative to the adjacent part of the static zone. When the inner side wall includes a concave portion, it is preferable that such a concave portion forms less than about 25, 10, 5, or 0.1% of the total outward facing surface area represented by the inner side wall. Preferably, the ratio of the total surface area of the internal structure that is in direct contact with the reaction medium to the total volume of the reaction zone is less than about 1, 0.5, 0.3 or 0.15 square meters per cubic meter. Preferably, the volume center of gravity of the static zone is horizontally offset from the volume center of gravity of the main reaction zone by less than about 0.4, 0.2, 0.1 or 0.01 parts of the maximum horizontal diameter of the main reaction zone.

When the bubble column reactor includes more than one internal structure defining more than one static zone, it is preferable that the static zones be arranged vertically in a row so that the volume center of gravity of all the static zones taken into account together is horizontally offset from the volume the center of gravity of the reaction zone is approximately less than 0.4, 0.2, 0.1 or 0.01 parts of the maximum horizontal diameter of the main reaction zone. Furthermore, when a plurality of static zones are formed inside the main reaction zone, it is preferable that the number of individual static zones having a volume of more than 0.2% of the volume of the main reaction zone is less than about 100, 10, 5, or 2.

The external reaction vessel of the bubble column reactor preferably has a ratio of maximum vertical height to maximum horizontal diameter in the range of about 3: 1 to 30: 1, more preferably in the range of about 6: 1 to 20: 1, and most preferably in the range of 9 : 1 to 15: 1. The internal structure preferably has a ratio of maximum vertical height to maximum horizontal diameter in the range of about 0.3: 1 to 100: 1, more preferably in the range of about 1: 1 to 50: 1, and most preferably in the range of 3: 1 to 30 :one. Preferably, the maximum horizontal diameter of the internal structure is in the range of about 0.1 to 5 m, more preferably in the range of about 0.3 to 4 m, and most preferably in the range of 1 to 3 m. Preferably, the maximum vertical height of the internal structure is in the range of about 1 to 100 m, more preferably in the range of about 3 to 50 m, and most preferably in the range of 10 to 50 m. Preferably, the maximum horizontal diameter of the inner page uktury is in the range from about 5 to 80, more preferably from about 10 to 60, most preferably from 20 to 50% of the maximum horizontal diameter of the external reaction vessel. Preferably, the maximum vertical height of the internal structure 602 is in the range of about 3 to 100% of the maximum vertical height of the external reaction vessel, more preferably in the range of about 10 to 90% of the maximum vertical height of the external reaction vessel, and most preferably in the range of 30 to 80 % of the maximum vertical height of the external reaction vessel. Any parameters (e.g., height, width, area, volume, relative horizontal placement and relative vertical placement) defined here for the external reaction vessel and structural parts are also considered applicable to the reaction zone defined by the external reaction vessel, and vice versa.

In one embodiment, the internal structure completely isolates the static zone from the reaction zone. In another embodiment of the invention, the internal structure defines one or more straight holes, which allows direct liquid transfer between the static zone and the reaction zone. When the internal structure defines such straight holes, it is preferable that the maximum diameter of the smallest of the straight holes is less than about 0.3, 0.2, 0.1, or 0.05 parts of the maximum horizontal diameter of the main reaction zone. When the internal structure defines such straight holes, it is preferable that the maximum diameter of the largest of the straight holes is less than about 0.4, 0.3, 0.2, or 0.1 parts of the maximum horizontal diameter of the main reaction zone. When the internal structure defines such straight holes, it is preferable that the total cross-sectional area determined by all straight holes is less than about 0.4, 0.3, or 0.2 parts of the maximum horizontal cross-sectional area of the main reaction zone. The internal structure has a maximum height (H i ). When the internal structure defines one or several lines of holes, preferably to less than about 50, 25 or 10% of the total area of the cross section crossing defined by all straight holes are at a distance more than about 0.5H i, or i 0.25H 0 , 1H i from the top of the internal structure. When a plurality of internal structures are used in a bubble column reactor to form a plurality of static zones, it is permissible for two or more static zones to include interconnected openings and / or pipelines that allow fluid to pass between the static zones. Preferably, the maximum diameter of the smallest of any interconnected openings and / or pipelines is less than about 0.3, 0.2, 0.1, or 0.05 parts of the maximum horizontal diameter of the main reaction zone.

As mentioned above, some of the physical and operational characteristics of the bubble column reactors described above using FIG. 1-25, create vertical gradients of pressure, temperature and concentrations of reagents (i.e. oxygen and an oxidizable compound) of the processed reaction medium. As discussed above, such vertical gradients can provide a more efficient and economical oxidation process than conventional oxidation processes, which prefer a well-mixed reaction medium with relatively uniform pressure, temperature, and concentration of reactants throughout. The vertical gradients of oxygen capable of oxidizing a compound (e.g., para-xylene) and temperature, which make it possible to use the oxidation system in accordance with an embodiment of the present invention, will be described in more detail.

As for FIG. 26, in order to quantify the concentration gradients of the reactants in the reaction medium during oxidation in the reactor as a bubble column, the entire volume of the reaction medium can be theoretically distributed into 30 discrete horizontal thin layers of equal volume. FIG. 26 illustrates the principle of dividing the reaction medium into 30 discrete horizontal thin layers of equal volume. With the exception of the uppermost and lowest horizontal thin layers, each horizontal thin layer is a discrete volume bounded at the top and bottom by imaginary horizontal planes and bounded on its sides by the wall of the reactor. The uppermost horizontal thin layer is bounded below by an imaginary horizontal plane, and above by the upper surface of the reaction medium. The lowest horizontal thin layer is bounded at the top by an imaginary horizontal plane, and below the bottom of the tank shell. After the distribution of the reaction medium into 30 discrete horizontal thin layers of equal volume, one can determine the time-averaged and volume-averaged concentration of each horizontal thin layer. A single horizontal thin layer having a maximum concentration of all 30 horizontal thin layers can be identified as “C-max horizontal thin layer”. A single horizontal thin layer located above the C-max horizontal layer and having a minimum concentration of all horizontal thin layers located above the C-max horizontal thin layer can be identified as “C-min horizontal thin layer”. The vertical concentration gradient can then be calculated as the ratio of the concentration in the C-max horizontal thin layer to the concentration in the C-min horizontal thin layer.

As for the quantitative determination of the oxygen concentration gradient, when the reaction medium is theoretically distributed into 30 discrete horizontal thin layers of equal volume, the O 2 -max horizontal thin layer is defined as the layer having the maximum oxygen concentration of all 30 horizontal thin layers, and O 2 -min horizontal a thin layer is defined as a layer having a minimum oxygen concentration of horizontal thin layers located above the O 2 -max of the horizontal thin layer. The oxygen concentration in the horizontal thin layers is measured in the gas phase of the reaction medium, based on the wet averaged over time and averaged over the volume of the molar composition. Preferably, the ratio of the oxygen concentration of the O 2 -max of the horizontal thin layer to the oxygen concentration of the O 2 -min of the horizontal thin layer is in the range of about 2: 1 to 25: 1, more preferably in the range of about 3: 1 to 15: 1, and most preferably in the range of 4: 1 to 10: 1.

Typically, the O 2 -max horizontal thin layer will be located near the lower part of the reaction medium, while the O 2 -min horizontal thin layer will be located near the upper part of the reaction medium. Preferably, the O 2 -min horizontal thin layer is one of the 5 uppermost horizontal thin layers of 30 discrete horizontal thin layers. Most preferably, the O 2 -min horizontal thin layer is one of the uppermost layers of 30 discrete horizontal thin layers, as shown in FIG. 26. Preferably, the O 2 -max horizontal thin layer is one of the 10 lowest horizontal thin layers of 30 discrete horizontal thin layers. Most preferably, the O 2 max horizontal thin layer is one of the 5 lowest horizontal thin layers of 30 discrete horizontal thin layers. For example, in FIG. 26 shows an O 2 -max horizontal thin layer in the form of a third horizontal thin layer from the bottom of the reactor. Preferably, the vertical gap between the O 2 min and O 2 max horizontal thin layers is at least about 2W, more preferably at least about 4W and most preferably at least 6W. Preferably, the vertical gap between the O 2 min and O 2 max horizontal thin layers is at least about 0.2 N, more preferably at least about 0.4 N and most preferably at least 0, 6H.

The time-averaged and volume-averaged oxygen concentration, based on wet oxygen, the O 2 -min horizontal thin layer is preferably in the range of about 0.1 to 3 mol%, more preferably in the range of about 0.3 to 2% pier and most preferably in the range from 0.5 to 1.5 mol%. The time-averaged and volume-averaged oxygen concentration, based on wet oxygen, the O 2 -max of the horizontal thin layer is preferably in the range of about 4 to 20 mol%, more preferably in the range of about 5 to 15 mol%. and most preferably in the range from 6 to 12 mol%. The time-averaged oxygen concentration, based on dry oxygen, in the gaseous effluent discharged from the reactor through the gas outlet is preferably in the range of about 0.5 to 9 mol%, more preferably in the range of 1 to 7 mol% . and most preferably in the range from 1.5 to 5 mol%.

Since the oxygen concentration drops so significantly towards the top of the reaction medium, it is preferable that the oxygen consumption is reduced in the upper part of the reaction medium. Such reduced consumption near the top of the reaction medium can be achieved by creating a vertical concentration gradient of the oxidizable compound (e.g., para-xylene), where the minimum concentration of the oxidizable compound is near the top of the reaction medium.

As for the quantitative determination of the concentration gradient of an oxidizable compound (for example, para-xylene), the reaction medium is theoretically distributed into 30 discrete horizontal thin layers of equal volume, the OS-max horizontal thin layer is identified as the layer having the maximum concentration of oxidizable compound from all 30 horizontal layers, and the OS-min horizontal thin layer is identified as a layer having a minimum concentration of oxidizable compounds from horizontal layers located above the OC-max horizontal thin layer. The concentrations of the oxidizable compound in thin horizontal layers are measured in the liquid phase based on a time-averaged and volume-averaged mass fraction. Preferably, the ratio of the concentration of the oxidizable compound OS-max horizontal thin layer to the concentration of the oxidizable compound OS-min horizontal thin layer is more than about 5: 1, more preferably more than about 10: 1, even more preferably more than about 20: 1 and most preferably ranged from 40: 1 to 1000: 1.

Typically, the OS-max horizontal thin layer will be located near the lower part of the reaction medium, while the OS-min horizontal thin layer will be located near the upper part of the reaction medium. Preferably, the OS-min horizontal thin layer is one of the 5 uppermost horizontal thin layers of 30 discrete horizontal thin layers. Most preferably, the OS-min horizontal thin layer is the uppermost layer of 30 discrete horizontal thin layers, as shown in FIG. 26. Preferably, the OS-max horizontal thin layer is one of the 10 lowest horizontal thin layers of 30 discrete horizontal thin layers. Most preferably, the OS-max horizontal thin layer is one of the 5 lowest horizontal thin layers of 30 discrete horizontal thin layers. For example, in FIG. 26 shows an OS-max horizontal thin layer in the form of a fifth horizontal thin layer from the reactor bottom. Preferably, the vertical gap between OC-min and OC-max of horizontal thin layers is at least about 2W, where W is the maximum width of the reaction medium. More preferably, the vertical gap between OC-min and OC-max of horizontal thin layers is at least about 4W and most preferably at least 6W. At a given height H of the reaction medium, it is preferred that the vertical gap between OC-min and OC-max of horizontal thin layers is at least about 0.2 N, more preferably at least about 0.4 N and most preferably at least least 0.6N.

The time-averaged and volume-averaged concentration of an oxidizable compound (e.g., para-xylene) in the OS-min liquid phase of the horizontal thin layer is preferably less than about 5,000 ppm, more preferably less than about 2,000 ppm. ./mln, even more preferably less than about 400 parts by weight per million, and most preferably is in the range from 1 to 100 parts by weight per million. The time-averaged and volume-averaged concentration of the oxidizable compound in the liquid phase of the OS-max horizontal thin layer is preferably in the range of about 100 to 10,000 parts by weight, more preferably in the range of about 200 to 5,000 parts by weight. / million and most preferably in the range from 500 to 3000 parts by weight / million

Although it is preferred that the bubble column reactor generates vertical concentration gradients of the oxidizable compound, it is also preferred that the volume percent of the reaction medium having the oxidizable compound in the liquid phase is higher than 1000 ppm by weight. Preferably, the time-averaged volume percent of the reaction medium having a concentration of the oxidizable compound in the liquid phase of more than 1000 ppmw is less than about 9%, more preferably less than about 6%, and most preferably less than 3%. Preferably, the time-averaged volume percent of the reaction medium having a concentration of the oxidizable compound in the liquid phase of more than 2500 ppm is less than about 1.5%, more preferably less than about 1% and most preferably less than 0, 5%. Preferably, the time-averaged volume percent of the reaction medium having a concentration of the oxidizable compound in the liquid phase of more than 10000 ppm is less than about 0.3%, more preferably less than about 0.1% and most preferably less than 0.03%. Preferably, the time-averaged volume percent of the reaction medium having a concentration of the oxidizable compound in the liquid phase of more than 25,000 ppm is less than about 0.03%, more preferably less than about 0.015% and most preferably less than 0.007% . It has been found that it is not required that the volume of the reaction medium having elevated levels of the oxidizable compound be in one continuous volume. In many cases, the chaotic flow pattern in the reaction vessel as a bubble column gives simultaneously two or more solid, but isolated parts of the reaction medium having elevated levels of oxidizable compound. For each time used in time averaging, all such continuous but isolated volumes are more than 0.0001% vol. from the entire reaction medium are added together to determine the total volume having elevated levels of concentrations of the oxidizable compound in the liquid phase.

In addition to the gradients of oxygen concentrations and the oxidizable compound discussed above, it is preferred that a temperature gradient exist in the reaction medium. If you consider again FIG. 26, this temperature gradient can be quantified in a manner analogous to the method for determining concentration gradients by distributing the reaction medium into 30 discrete horizontal thin layers of equal volume and measuring the time-averaged and volume-averaged temperature of each thin layer. Then the horizontal thin layer with the lowest temperature of the lowest 15 horizontal thin layers can be identified as T-min horizontal thin layer, and the horizontal thin layer located above the T-min horizontal thin layer and having the maximum temperature of all thin layers above T- min horizontal thin layer, can be identified as "T-max horizontal thin layer". Preferably, the T-max temperature of the horizontal thin layer is at least 1 ° C higher than the T-min temperature of the horizontal thin layer. More preferably, the temperature T-max of the horizontal thin layer is in the range of about 1.25 to 12 ° C. higher than the temperature T-min of the horizontal thin layer. Most preferably, the T-max temperature of the horizontal thin layer is in the range of about 2 to 8 ° C. higher than the T-min temperature of the horizontal thin layer. The temperature T-max of the horizontal thin layer is preferably in the range of about 125 to 200 ° C, more preferably in the range of about 140 to 180 ° C, and most preferably in the range of 150 to 170 ° C.

Typically, the T-max horizontal thin layer will be located near the center of the reaction medium, while the T-min horizontal thin layer will be located near the bottom of the reaction medium. Preferably, the T-min horizontal thin layer is one of the 10 lowest horizontal thin layers of the 15 lowest horizontal thin layers. Most preferably, the T-min horizontal thin layer is one of the 5 lowest horizontal horizontal thin layers of the 15 lowest horizontal thin layers. For example, FIG. 26 shows the T-min horizontal thin layer as the second horizontal thin layer from the bottom of the reactor. Preferably, the T-max horizontal thin layer is one of 20 middle horizontal thin layers of 30 discrete horizontal thin layers. Most preferably, the T-min horizontal thin layer is one of 14 middle horizontal thin layers of 30 discrete horizontal thin layers. For example, in FIG. 26 shows a T-max horizontal thin layer in the form of the twentieth horizontal thin layer (i.e., one of 10 middle horizontal thin layers) from the bottom of the reactor. Preferably, the vertical gap between the T-min and T-max horizontal thin layers is at least about 2W, more preferably at least about 4W and most preferably at least 6W. Preferably, the vertical gap between the T-min and T-max horizontal thin layers is at least about 0.2 N, more preferably at least about 0.4 N and most preferably at least 0.6 N.

As described above, when there is a vertical temperature gradient in the reaction zone, it can provide the advantages of removing the reaction medium at a height-elevated location where the temperature of the reaction medium is highest, especially when the withdrawn product is subjected to further processing in the flow direction at higher temperatures. Therefore, when the reaction medium 36 is removed from the reaction zone through one or more elevated outlet openings, as shown in FIG. 15 and 16, it is preferred that the outlet (s) raised in height (s) be positioned (s) near the T-max of a horizontal thin layer. Preferably, the height-raised outlet is located within 10 horizontal thin layers of the T-max horizontal thin layer, more preferably within 5 horizontal thin layers of the T-max horizontal thin layer, and most preferably within 2 horizontal thin layers of the T-max horizontal thin layer.

It should be noted that many of the claimed features described in the present invention can be used in composite reactor oxidation systems - not only for systems that use a single oxidation reactor. In addition, some of the claimed features described in the present invention can be used in oxidation reactors with mechanical stirring and / or with stirring with a stream - not only in reactors using bubble mixing (i.e. in bubble column reactors). For example, several advantages are disclosed associated with a stepwise / variable oxygen concentration and / or oxygen consumption rate throughout the reaction medium. The advantages realized by the stepwise concentration / absorption of oxygen in the reaction medium can be achieved regardless of whether the entire volume of the reaction medium is contained in one vessel or in multiple vessels. In addition, the benefits realized by the stepwise concentration / consumption of oxygen in the reaction medium can be achieved regardless of whether the reaction vessel (s) is mechanically agitated (s), agitated (s) by flow and / or mixed (s) bubbles.

One method for quantifying the stepwise change in oxygen concentration and / or oxygen consumption rate in a reaction medium is to compare two or more discrete 20% continuous volumes of the reaction medium. Such 20% discrete continuous volumes do not require any particular form of limitation. However, each 20% continuous volume must be formed by a continuous volume of the reaction medium (that is, each volume is “continuous”), and the 20% continuous volumes must not overlap with each other (that is, the volumes are “discrete”). Such discrete 20% continuous volumes may be located in the same reactor (FIG. 29) or in multiple reactors. In FIG. 27, a bubble column reactor is shown as containing a reaction medium that includes a first discrete 20% continuous volume 37 and a second discrete 20% continuous volume 39.

A stepwise change in the availability of oxygen in the reaction medium can be quantified with reference to a 20% continuous volume of the reaction medium having the most enriched molar fraction of oxygen in the gas phase, and with reference to a 20% continuous volume of the reaction phase having the most depleted molar fraction of oxygen in the gas phase. In the gas phase of a discrete 20% continuous volume of the reaction medium containing the highest oxygen concentration in the gas phase, the time-averaged and volume-averaged oxygen concentration based on wet oxygen is preferably in the range of about 3 to 18 mol%, more preferably in the range of from about 3.5 to 14 mol%. and most preferably in the range from 4 to 10 mol%. In the gas phase of a discrete 20% continuous volume of the reaction medium containing the lowest concentration of oxygen in the gas phase, time averaged and volume averaged oxygen concentration based on wet oxygen, is preferably in the range of about 0.3 to 5 mol% ., more preferably in the range of from about 0.6 to 4 mol%. and most preferably in the range from 0.9 to 3 mol%. In addition, the ratio of the time-averaged and volume-averaged oxygen concentrations based on wet oxygen in the most enriched 20% continuous volume of the reaction medium relative to the most depleted 20% continuous volume of the reaction medium is preferably in the range of about 1.5: 1 to 20: 1, more preferably in the range of about 2: 1 to 12: 1, and most preferably in the range of 3: 1 to 9: 1.

A step change in the rate of oxygen consumption in the reaction medium can be quantified in terms of oxygen SOS (oxygen-STR), as described previously. The concept of “oxygen SOS” was previously described in a general sense (that is, from a forecast of the average oxygen SOS of the entire reaction medium); however, the oxygen SOS can also be considered in the local sense (i.e., for a part of the reaction medium) in order to quantify the stepwise change in oxygen consumption throughout the reaction medium.

It has been found that it is very useful to induce a change in the oxygen concentration of the whole reaction medium in general agreement with the desired gradients disclosed in the present invention relating to the pressure in the reaction medium and the molar fraction of molecular oxygen in the gas phase of the reaction medium. Therefore, it is preferable that the ratio of the oxygen SOS of the first discrete 20% continuous volume of the reaction medium to the oxygen SOS of the second discrete 20% continuous volume of the reaction medium is in the range from about 1.5: 1 to 20: 1, more preferably in the range from about 2: 1 to 12: 1, and most preferably in the range from 3: 1 to 9: 1. In one embodiment of the invention, the “first discrete 20% continuous volume of the reaction medium” is closer than the “second discrete 20% continuous volume of the reaction medium” to the location where molecular oxygen is initially introduced into the reaction medium. Such large gradients of oxygen SOC are desirable, regardless of whether the partial oxidation reaction medium is in the oxidation reactor as a bubble column or in any other type of reaction vessel, where pressure and / or molar fractions of molecular oxygen in the gas phase in the reaction medium are created ( for example, in a vessel with mechanical stirring, having many vertically arranged mixing zones, achieved through the use of multiple blades having a strong radial flow, possibly up to filled with horizontal deflector devices, with an oxygen flow rising, as a rule, upward from the feed near the lower part of the reaction vessel, despite the fact that a significant back-mixing of the oxidizing stream can take place within each vertically located mixing zone and that some back-mixing oxidant flow may occur between adjacent vertically adjacent mixing zones). That is, when there is a gradient of pressure and / or molar fraction of molecular oxygen in the gas phase of the reaction medium, the applicants have found that it is desirable to create a similar gradient of chemical consumption of dissolved oxygen using the means disclosed in the invention.

A preferred means of causing a local measurement of oxygen SOS is to control the feed locations of the oxidizable compound and to control the mixing of the liquid phase of the reaction medium in order to control the concentration gradients of the oxidizable compound in accordance with the description of the present invention. Another useful tool that causes a local change in oxygen SOS is to change the reaction activity by creating a local temperature fluctuation and by changing the local mixture of catalyst and solvent components (for example, by introducing additional gas to cause cooling by evaporation in a certain part of the reaction medium; and by adding a solvent stream containing a higher amount of water to reduce activity in a particular part of the reaction medium).

When the oxidation reactor has the configuration "reactor in the reactor", which is described above with reference to FIG. 12-14, it is preferable that the concentration gradients, temperature gradients and oxygen COS gradients described in the present invention with reference to FIG. 26 and 27 were applicable to a portion of the reaction medium located inside the external reactor and outside the internal reactor (for example, to the reaction medium 220a of FIG. 12).

As for FIG. 1-27, the oxidation is preferably carried out in a bubble column reactor under conditions that are very different in accordance with the preferred embodiments described in the present invention from the conditions of conventional oxidation reactors. When using a bubble column reactor for conducting liquid-phase partial oxidation of para-xylene to crude terephthalic acid (CTK) in accordance with the preferred embodiments described in the invention, spatial profiles of local reaction intensity, local evaporation rate, and local temperature combined with flow patterns liquids within the reaction zone, and preferred relatively low oxidation temperatures contribute to the formation of STK particles having unique and useful properties.

In FIG. 28A and 28B illustrate STK base particles produced in accordance with one embodiment of the present invention. In FIG. 28A shows STK base particles at a magnification of 500 times, while FIG. 28B details the image on one of the STK base particles and shows this particle at 2000 times magnification. As, apparently, the most well represented in FIG. 28B, each STK base particle is typically composed of a large number of small, agglomerated STK subparticles, resulting in a STK base particle with relatively high surface area, high porosity, low density and good solubility. Unless otherwise specified, the various properties of the inventive STK described below are measured using a typical STK sample, where a typical sample weighs at least 1 gram and / or is formed from at least 10,000 individual STK particles. STK base particles typically have an average particle size in the range of about 20 to 150 microns, more preferably in the range of about 30 to 120 microns, and most preferably in the range of 40 to 90 microns. STC subparticles typically have an average particle size in the range of about 0.5 to 30 microns, more preferably in the range of about 1 to 15 microns, and most preferably in the range of 2 to 5 microns. The relatively high surface area of the base particles of the STC shown in FIG. 28A and 28B can be quantified using the Braunauer-Emmett-Teller (BET) method of measuring surface area. Preferably, the STK base particles have an average BET surface area of at least about 0.6 square meters per gram (m 2 / g). More preferably, the STK base particles have an average BET surface area in the range of about 0.8 to 4 m 2 / g. Most preferably, STK base particles have an average BET surface area in the range of about 0.9 to 2 m 2 / g. Physical properties (for example, particle size, BET surface area, porosity and solubility) of STK base particles obtained using the optimized oxidation process of the preferred embodiment of the present invention allow STK particles to be cleaned using more efficient and / or economical methods that are described hereinafter in more detail with reference to FIG. 31.

The mean particle sizes given above are determined using polarized light microscopy and image analysis. The equipment used for particle size analysis includes a Nikon E800 optical microscope with a 4x Plan Flour NA 0.13 lens, a Spot RT TM digital camera, and a personal computer with Image Pro Plus TM V4.5.0.19 image analysis software. The method of particle size analysis includes the following main stages: (1) dispersing STK powders in mineral oil; (2) preparation of a slide / coverslip system for dispersion; (3) the study of a glass slide using a microscope in polarized light (conditions of intersecting polar - particles appear in the form of bright objects on a dark background); (4) fixing different images for each sample of the preparation (field size = 3 × 2.25 mm; pixel size = 1.84 μm / pixel); (5) carrying the image analysis using the software Image Pro Plus TM; (6) transferring particle measurements to a pivot table; and (7) carrying out a statistical description in a pivot table. Stage (5) “conducting image analysis using Image Pro Plus TM software” includes the following steps: (a) setting an image threshold to detect white particles on a black background; (b) creating a binary image; (c) passing a single pass open filter to remove pixel noise; (d) measuring all particles in the image; and (e) fixing the average diameter measured for each particle. Software Image Pro Plus TM defines mean diameter of individual particles as the number average length of diameters of the particles measured at 2 degree intervals and passing through the center of gravity of the particle. Stage 7 "conducting a statistical description in the summary table" includes the calculation of the volume-weighted average particle size as follows. The volume of each of the n particles in the sample is calculated as if it were spherical using pi / 6 * d i ^ 3; multiply the volume of each particle by its diameter to find pi / 6 * d i ^ 4; summarize for all particles in the sample the values pi / 6 * d i ^ 4; summarize the volumes of all particles in the sample and calculate the volume-weighted particle diameter as the sum for all n particles in the sample (pi / 6 * d i ^ 4) divided by the sum of all n particles in the sample (pi / 6 * d i ^ 3) . As used herein, “average particle size” refers to a volume-weighted average particle size determined in accordance with the test method described above; it is also called as D (4,3).

Figure 00000001

In addition, stage (7) involves the identification of particle sizes for which various fractions of the entire sample volume are smaller. For example, D (v, 0.1) is the particle size for which 10% of the total sample volume is smaller and 90% are larger; D (v, 0.5) is the particle size for which one half of the total volume of the sample is larger and the other half is smaller; D (v, 0.9) is the particle size for which 90% of the total sample volume is smaller; etc. In addition, stage (7) involves calculating the value of D (v, 0.9) minus D (v, 0.1), which in this case is defined as the “particle size dispersion”; and step (7) involves calculating the size dispersion of the particles divided by D (4.3), which is defined in this case as “relative particle size dispersion”.

In addition, it is preferable that the D (v, 0.1) STK particles measured as described above are in the range of about 5 to 65 microns, more preferably in the range of about 15 to 55 microns, and most preferably in the range of 25 up to 45 microns. Preferably, the D (v, 0.5) STK particles measured as described above are in the range of about 10 to 90 microns, more preferably in the range of about 20 to 80 microns, and most preferably in the range of 30 to 70 microns . Preferably, the D (v, 0.9) STK particles, measured as described above, are in the range of about 30 to 150 microns, more preferably in the range of about 40 to 130 microns, and most preferably in the range of 50 to 110 microns . Preferably, the relative particle size dispersion is in the range of about 0.5 to 2.0, more preferably in the range of about 0.6 to 1.5, and most preferably in the range of 0.7 to 1.3.

The BET surface area values presented above were measured on a Micromeritics ASAP2000 instrument (Micromeritics Instrument Corporation of Norcross, GA). At the first stage of the measurement process, 2-4 grams of a sample of particles are weighed and dried in vacuum at 50 ° C. Then the sample is placed on the collector for gas analysis and cooled to 77ºK. The nitrogen adsorption isotherm is measured at 5 minimum equilibrium pressures by exposing the sample to known volumes of nitrogen and measuring the pressure drop. The equilibrium pressures are approximately in the range P / P o = 0.01-0.20, where P is the equilibrium pressure and P o is the vapor pressure of liquid nitrogen at 77K. The resulting isotherm is then plotted in accordance with the following BET equation:

Figure 00000002

where V a is the volume of gas adsorbed by the sample at P, V m is the volume of gas required to cover the entire surface of the sample with one layer of gas, and C is a constant. From the obtained graph, V m and C are determined. Then V m is transferred to the surface area using the nitrogen cross-sectional area at 77K using the expression:

Figure 00000003

where σ is the nitrogen cross-sectional area at 77K, T is 77K, and R is the gas constant.

As mentioned above, STK obtained in accordance with one of the variants of the present invention, shows excellent solubility characteristics relative to STK obtained by other processes. This increased degree of dissolution allows the cleaning of the inventive STK to be carried out more efficiently and / or using more efficient cleaning processes. The following description relates to a method by which STK dissolution rate can be quantified.

The dissolution rate of a known amount of solids in a known amount of solvent while mixing the mixture can be determined by various methods. The method used in this case, called the “timed dissolution test”, is described below. An ambient pressure of approximately 0.1 MPa is used throughout the entire timed dissolution test. An ambient temperature of approximately 22 ° C is used throughout the entire timed dissolution test. In addition, solids, solvent and all dissolution equipment are fully thermally balanced at this temperature prior to the start of the test, and during dissolution there is no noticeable heating or cooling of the beaker and its contents. A portion of a solvent, fresh analytical grade tetrahydrofuran for HPLC (purity> 99.9), then a 250 g THF, were placed in a 400 ml KIMAX clean tall glass beaker (Kimble ® lot number 14020, Kimble / Kontes, Vineland, NJ ), which is non-insulated, smooth-walled and has a substantially cylindrical shape. A Teflon-coated magnetic stirrer (VWR lot number 58948-230, approximately 1 inch long and 3/8 inch diameter with an octagonal cross section, VWR International, West Chester, PA 19380) is placed in a beaker where it naturally settles to the bottom. The sample was mixed using a 15-point multi-point Variomag® magnetic stirrer (H&P Labortechnik AG, Oberschleissheim, Germany) at 800 rpm. This mixing is started no more than 5 minutes before the addition of solids and continues continuously for at least 30 minutes after adding solids. A solid sample of crude or purified TFA in the form of particles in an amount of 250 mg is weighed in a weight cup to which the sample does not adhere. At the initial time, denoted as t = 0, suspended solids contribute all in one portion to the mixed THF and at the same time start the timer. Essentially, THF wets solids very quickly and forms a diluted, well-mixed suspension within 5 seconds. Then samples of this mixture are taken at the following time intervals, measured in minutes from t = 0: 0.08, 0.25, 0.50, 0.75, 1.00, 1.50, 2.00, 2.50, 3.00, 4.00, 5.00, 6.00, 8.00, 10.00, 15.00 and 30.00. Each small sample was taken from a diluted, well-mixed mixture using a new disposable syringe (Becton, Dickinson and Co., 5 ml, REF 30163, Franklin Lakes, NJ 07417). Immediately after sampling from a beaker, approximately 2 ml of a clear liquid sample is quickly discharged through a new, unused syringe filter (25 mm diameter, 0.45 μm, Gelman GHP Acrodisc GF ® , Pall Corporation, East Hills, NY 11548) into a new labeled glass vessel for samples. The duration of each filling of the syringe, the installation of the filter and the discharge into the sample vessel is less than about 5 seconds, and this interval is appropriate to start and end within about 3 seconds from either side of each selected sampling time. Within approximately five minutes from each filling, the sample vessels are lidded and stored at approximately constant temperature until subsequent chemical analyzes. After the last sample was taken 30 minutes after t = 0, all six samples were analyzed for the amount of dissolved TFA using the HPLC-DAD method, which is generally described later in this description. However, in the test under consideration, the calibration standards and the results presented in both cases are based on milligrams of dissolved TFA per gram of solvent, THF (hereinafter “ppm in THF). For example, if all 250 mg of solids are very pure TFA and if all of this is completely dissolved in 250 g of THF before the selection of the appropriate sample, the exact measured concentration will be 1000 ppm in THF.

When the STC in accordance with the present invention is subjected to the timed dissolution test described above, it is preferable that a sample taken one minute after t = 0 dissolves to a concentration of at least about 500 ppm in THF, more than preferably at least 600 ppm in THF. For a sample taken two minutes after t = 0, it is preferable that the STK in accordance with the present invention is dissolved to a concentration of at least about 700 ppm in THF, more preferably at least 750 ppm in THF. For a sample taken four minutes after t = 0, it is preferable that the STK in accordance with the present invention is dissolved to a concentration of at least about 840 ppm in THF, more preferably at least 880 ppm THF

It has been established that a relatively simple development model with a negative exponent can be used to describe the time dependence of the entire data set for a complete test for timed dissolution, despite the complexity of the samples in the form of particles and the dissolution process. The form of the equation, hereinafter referred to as the "model calculated by the time of dissolution", is as follows:

S = A + B * (1-exp (-C * t)),

where t is the time in minutes;

S is the solubility in ppm in THF at time t;

exp is the exponential function at the base of the natural logarithm of 2;

A, B are constant constants in ppm in THF, where A relates mainly to the rapid dissolution of smaller particles in a very short time and where the sum of A + B relates mainly to the total amount of solute at the end of a certain test period; and

C is the constant time constant in the return minutes.

The constant constants are adjusted to minimize the sum of the squared errors between the points of the actual data and the corresponding model values, and this method is usually called the “least squares” fit. The preferred software package for this data regression is JMP Release 5.1.2 (SAS Institute Inc., JMP Software, SAS Campus Drive, Cary NC 27513).

When an STK according to the present invention is tested using a timed dissolution test and approximated to a timed dissolution model as described above, it is preferable that the STK has a time constant “C” of more than about 0.5 reverse minutes, more preferably more than about 0.6 reverse minutes and most preferably more than 0.7 reverse minutes.

In FIG. 29A and 29B show a particle of a conventional STC obtained using a conventional high-temperature oxidation process in a continuous stirred reactor (LRC). In FIG. 29A shows a particle of conventional STK at 500x magnification, whereas in FIG. 29B, the image is enlarged, and the STK particle is shown at 2000 times magnification. A visual comparison of the particles of the claimed STK presented in FIG. 28A and 29B, and particles of conventional STC shown in FIG. 29A and 29B shows that a conventional STK particle has a higher density, lower surface area, lower porosity and larger particle size than the particles of the claimed STK. Indeed, the conventional STC presented in FIG. 29A and 29B, has an average particle size of approximately 205 μm and a BET surface area of approximately 0.57 m 2 / g.

FIG. 30 illustrates a typical manufacturing process for purified terephthalic acid (OTC). In a typical OTC process, para-xylene is partially oxidized in a high-temperature oxidation reactor 700 with mechanical stirring. The suspension containing the STK is withdrawn from the reactor 700 and then cleaned in the purification system 702. The product, OTC, purification systems 702 are introduced into the separation system 706 to separate and dry the OTC particles. Purification system 702 contributes a significant part to the costs associated with the production of TCI particles by conventional methods. The purification system 702 typically includes a water / exchange system 708, a dissolution system 710, a hydrogenation system 712, and three separate crystallization vessels 704a, b, c. In the 708 water / exchange system, a significant portion of the mother liquor is replaced with water. After adding water, the water / STK suspension is introduced into the dissolution system 710, where the water / STK mixture is heated until the STK particles are completely dissolved in water. After the STC is dissolved, the STK solution in water is hydrogenated in the hydrogenation system 712. The hydrogenated effluent from the hydrogenation system 712 is then subjected to three crystallization stages in crystallization vessels 704a, b, c, after which OTK is separated in the separation system 706.

FIG. 31 illustrates an improved OTC manufacturing process that utilizes a reactor oxidation system including a primary oxidation reactor 800a and a secondary oxidation reactor 800b. In the configuration presented in FIG. 31, an initial slurry is obtained from the primary oxidation reactor 800a, and then subjected to purification in a purification system 802, in which the secondary oxidation reactor 800b is a part. The initial suspension withdrawn from the primary oxidation reactor 800a preferably contains solid STK particles and a liquid mother liquor. Typically, the initial suspension contains in the range from about 10 to 50% of the mass. particles of solid STK, the rest being liquid mother liquor. Particulate solid CTK particles present in the initial slurry withdrawn from the primary oxidation reactor 800a typically contain at least about 400 parts by weight of 4-carboxybenzaldehyde (4-KBA), more typically at least about 800 ppm of 4-KBA and most typically in the range of about 1000 to 15000 ppm of 4-KBA.

Purification system 802 receives an initial slurry withdrawn from the primary oxidation reactor 800a and reduces the concentration of 4-KBA and other impurities present in the STK. In a purification system 802, a cleaner / purified suspension is obtained, and this suspension is separated and dried in a separation system 804, resulting in particles of a purer terephthalic acid solid containing less than about 400 ppm 4-KBA, more preferably less than about 250 parts by weight per million of 4-KBA, and most preferably in the range of about 10 to 200 parts by weight of 4-KBA.

Purification system 802 includes a secondary oxidation reactor 800b, a liquid exchange system 806, an autoclave 808, and one crystallizer 810. In the secondary oxidation reactor 800b, the initial suspension is oxidized at a temperature and pressure that are approximately equal to the temperature and pressure in the primary oxidation reactor 800a. In a fluid exchange system 806, at least 50% of the mass. the mother liquor present in the suspension withdrawn from the secondary oxidation reactor 800b is replaced with a fresh solvent, so as to result in a solvent suspension containing STC particles and a solvent. A suspension of the exchanged solvent leaving the fluid exchange system 806 is introduced into the autoclave 808. An additional oxidation reaction is carried out in the autoclave 808 at slightly higher temperatures than the temperatures used in the primary oxidation reactor 800a.

As described above, the high surface area, small particle size, and low density of STK particles obtained in the primary oxidation reactor 800a make some impurities trapped in STK particles available for oxidation in an autoclave 808 without the need for complete dissolution of the STK particles in an autoclave 808. Therefore , the temperature in the autoclave 808 may be lower than in many similar processes in the art. The additional oxidation carried out in an autoclave 808 preferably reduces the concentration of 4-CBA in the STC by at least 200 parts by weight per million, more preferably by at least about 400 parts by weight per million, and most preferably the decrease is in the range from 600 to 6000 parts by weight per million. Preferably, the burning temperature in the autoclave 808 is at least about 10 ° C higher than the primary oxidation temperature in the reactor 800a, more preferably about 20-80 ° C higher than the primary oxidation temperature in the reactor 800a, and most preferably 30-50 ° C higher than primary oxidation temperature in the reactor 800a. The combustion temperature is preferably in the range of about 160 to 240 ° C, more preferably in the range of about 180 to 220 ° C, and most preferably in the range of about 190 to 210 ° C. For a purified product from an autoclave 808, only one crystallization step is required in the crystallizer 810 before being isolated in the separation system 804. Suitable secondary oxidation / incineration procedures are described in more detail in US Patent Application No. 2005/0065373, the entire disclosure of which is hereby incorporated by reference.

Terephthalic acid (e.g., OTC) produced using the system of FIG. 31 preferably forms particles having an average particle size of at least about 40 microns, more preferably in the range of about 50 to 2000 microns, most preferably in the range of about 60 to 200 microns. The TCA particles preferably have an average BET surface area of less than about 0.25 m 2 / g, more preferably in the range of about 0.005 to 0.2 m 2 / g, and most preferably in the range of 0.01 to 0.18 m 2 / g OTC obtained using the system presented in FIG. 31 is acceptable for use as a raw material for the production of PET. Typically, PET is prepared by esterification of terephthalic acid with ethylene glycol followed by polycondensation. Preferably, terephthalic acid produced using embodiments of the present invention is used as a feedstock for the PET production process in a tubular reactor described in US Patent Application Serial No. 10/013318 for review on December 7, 2001, the entire disclosure of which is incorporated by reference .

Particles of OTC with a preferred morphology described in this invention, especially can be used in the above oxidative combustion process to reduce the content of 4-KBA. In addition, such preferred STC particles provide advantages over a wide range of other subsequent processes in which dissolution and / or chemical interaction of the particles takes place. Such further post-treatments include, but are not limited to, reacting the at least one hydroxyl-containing compound to form ester compounds, in particular the reaction of STK with methanol to form dimethyl terephthalate and impurity esters; a reaction with at least one diol to form ester monomeric and / or polymeric compounds, in particular a reaction of STK with ethylene glycol to form polyethylene terephthalate (PET); and complete or partial dissolution in solvents, including, but not limited to, water, acetic acid and N-methyl-2-pyrrolidone, which may include additional processing, including, but not limited to, re-precipitation of purer terephthalic acid and / or a selective chemical reaction of carbonyl groups other than carboxylic acid groups. This is especially true for the significant dissolution of STK in a solvent containing water, combined with partial hydrogenation, which reduces the amount of aldehydes, especially 4-KBA, fluorenones, phenons and / or anthraquinones.

In accordance with one embodiments of the present invention, there is provided a method for partial oxidation of an oxidizable aromatic compound to one or more types of aromatic carboxylic acids, wherein the purity of a portion of the solvent in the feed (ie, “solvent charge”) and the purity of the portion of the oxidizable compound in the feed (i.e., “loading of the oxidizable compound”) is controlled within certain ranges specifically defined below. Along with other embodiments of the present invention, this makes it possible to control the purity of the liquid phase and, if present, the solid phase and the combined suspension phase (i.e., solid plus liquid) of the reaction medium within certain preferred ranges described below.

Regarding the loading of the solvent, it is known to oxidize the oxidizable aromatic compound (s) (s) to produce aromatic carboxylic acid, where the loading of the solvent introduced into the reaction medium is a mixture of analytically pure acetic acid and water, as is often the case used in laboratory conditions and in pilot plants. Similarly, it is known to carry out the oxidation of an oxidizable aromatic compound to an aromatic carboxylic acid, where the solvent leaving the reaction medium is separated from the produced aromatic carboxylic acid and then recycled back to the reaction medium as solvent charge, mainly due to production costs. Such recycling of the solvent over time leads to the accumulation of some impurities in the feed and process by-products in the recycled solvent. Various means are known in the art that help purify the recycled solvent before reintroducing it into the reaction medium. Typically, a higher degree of purification of the recycled solvent leads to a significantly higher production cost than in the case of a lower degree of purification using similar means. One of the embodiments of the present invention relates to the interpretation and determination of preferred intervals for a large number of impurities in the loading of the solvent, many of which were previously considered to be largely acceptable in order to find the optimal balance between the total cost of production and the overall purity of the product.

“Recycled solvent charge” is defined herein as the charge of a solvent that previously formed part of the reaction medium that has been oxidized in the oxidation zone / reactor and has left the oxidation zone / reactor as part of a crude liquid and / or suspension product. For example, loading a recycled solvent into a partial oxidation reaction medium during the oxidation of para-xylene to form TFA is a solvent that initially formed part of the partial oxidation reaction medium, was removed from the reaction medium as a liquid phase of a suspension of TFA, separated from most of the solid TFA and then returned to the partial oxidation reaction medium. As described above, such a recycled solvent charge tends to accumulate all kinds of undesirable impurities unless specific additional process steps are provided for purifying the solvent at significant capital and operating costs. For economic reasons, it is preferable that at least about 20% of the mass. loading the solvent into the reaction medium of the present invention was a recycled solvent, more preferably at least about 40% by weight, even more preferably at least about 80% by weight. and most preferably at least about 90% of the mass. Based on the supply of solvent and the duration of the working cycle in an industrial plant, it is preferred that portions of the recycled solvent pass through the reaction medium at least once per day of operation, more preferably at least once a day for at least seven consecutive days of work, and most preferably at least once a day for at least 30 consecutive days of work.

Based on the reactivity and taking into account the metal impurities remaining in the oxidation product, it was found that the concentrations of the selected multivalent metals in the charge of the recycled solvent are preferably within the ranges specifically indicated immediately below. The concentration of iron in the recycled solvent is preferably below about 150 parts by weight per million, more preferably below about 40 parts by weight per million, and most preferably in the range from 0 to 8 parts by weight per million. The concentration of nickel in the recycled solvent is preferably below about 150 parts by weight per million, more preferably below about 40 parts by weight per million, and most preferably in the range from 0 to 8 parts by weight per million. The concentration of chromium in the recycled solvent is preferably below about 150 parts by weight per million, more preferably below about 40 parts by weight per million, and most preferably in the range from 0 to 8 parts by weight per million. The concentration of molybdenum in the recycled solvent is preferably below about 75 parts by weight / million, more preferably below about 20 parts by weight / million, and most preferably in the range from 0 to 4 parts by weight / million. The concentration of titanium in the recycled solvent is preferably below about 75 parts by weight / million, more preferably below about 20 parts by weight / million, and most preferably in the range from 0 to 4 parts by weight / million. The concentration of copper in the recycled solvent is preferably below about 20 parts by weight per million, more preferably below about 4 parts by weight per million, and most preferably in the range from 0 to 1 parts by weight per million. Other metallic impurities are typically also present in the recycled solvent, and usually their concentration varies at lower levels relative to one or more of the metals listed above. Controlling the above metals at preferred ranges will keep other metallic impurities at acceptable levels.

Such metals can appear as impurities in any feedstock entering the process (for example, in a feedable oxidizable compound, in a solvent, in an oxidizing agent, and in catalytic compounds). On the other hand, metals can appear as corrosion products of any of the process units in contact with the reaction medium and / or in contact with the recycled solvent. Means for controlling metals in the disclosed concentration ranges include appropriate descriptions and monitoring of the purity of various raw materials and the use of suitable structural materials, including, but not limited to, many industrial grades of titanium and stainless steel, including grades known as duplex stainless steels and high molybdenum stainless steels.

Applicants also disclose preferred ranges for selected aromatic compounds in a recycled solvent. These ranges include both precipitated and dissolved aromatic compounds in a recycled solvent.

Unexpectedly, even the precipitated product (e.g. TFA) of the partial oxidation of para-xylene is a pollutant that must be controlled in a recycled solvent. Since there are unexpectedly preferred ranges for solids levels in the reaction medium, any precipitated product in the solvent charge is immediately subtracted from the amount of oxidizable compound that can be fed together. Moreover, the supply of precipitated solid TFA in a recycled solvent at elevated levels was found to adversely affect the properties of particles formed in the oxidation medium with precipitation, which leads to undesirable properties in downflow operations (for example, product filtration, solvent washing, oxidative combustion of the crude product, dissolution of the crude product for subsequent processing, etc.). Another undesirable property of precipitated solids in the charge of the recycled solvent is that they often contain very high levels of precipitated impurities compared to the concentration of impurities in the volume of solids in TFA suspensions, from which most of the recycled solvent is obtained. Possibly, the increased levels of impurities observed in solids suspended in a recycled solvent may be related to nucleation time for precipitation of some impurities from the recycled solvent and / or cooling of the recycled solvent, either intentionally or due to external losses. For example, concentrations of intensely colored and undesired 2,6-dicarboxyfluorenone are observed at significantly higher levels in solids present in the recycled solvent at 80 ° C than in TFA solids separated from the recycled solvent at 160 ° C. Similarly, isophthalic acid concentrations are observed at significantly higher levels in solids present in the recycled solvent, compared to levels observed in TFA solids from the reaction medium. The way in which the particular precipitated impurities are involved in the recycled solvent upon re-introduction into the reaction medium, as it turns out, changes. This depends, apparently, on the relative solubility of the impurity in the liquid phase of the reaction medium, apparently, on how the deposited impurity is bedded within the deposited solids and, apparently, on the local deposition rate of TFA, where the solid is initially repeatedly enters the reaction medium. Thus, it was found useful to control the level of certain impurities in the recycled solvent, as described below, without considering whether these impurities are present in the recycled solvent in dissolved form or are involved in the material in the form of particles contained therein.

The amount of precipitated solids present in the recycled solvent is determined by the gravimetric method as follows. A typical sample is removed from the solvent supply to the reaction medium while the solvent flows through a conduit to the reaction medium. The useful size of the sample is approximately 100 g, taken in a glass container having approximately 250 ml of internal volume. Before bleeding to atmospheric pressure, but, with continuous flow in the direction of the sample container, the recycled solvent is cooled to less than 100 ° C; such cooling is carried out in order to limit the evaporation of the solvent for a short period before hermetically closing the glass container. After sampling at atmospheric pressure, the glass container is immediately sealed. The sample is then allowed to cool to approximately 20 ° C at an ambient temperature of approximately 20 ° C and without forced convention. Upon reaching 20 ° C, the sample is kept under such conditions for at least about 2 hours. The hermetically sealed container is then shaken vigorously until a uniform distribution of solids is achieved. Immediately afterwards, a magnetic stirrer is introduced into the sample container and rotated at a sufficient speed to effectively maintain a uniform distribution of solids. A 10 ml aliquot of the mixed liquid with suspended solids was pipetted and weighed. Then the volume of the liquid phase from this aliquot is separated by vacuum filtration, still at approximately 20 ° C and without loss of solids. Wet solids filtered from this aliquot are then dried without sublimating the solids, and such dried solids are weighed. The ratio of the mass of dried solids to the mass of the original aliquot of the suspension is a fraction of solids, usually expressed as a percentage, and in this case this fraction is called the amount of “solids deposited at 20 ° C” in the solvent charge.

Applicants have found that aromatic compounds dissolved in the liquid phase of the reaction medium and containing aromatic carboxylic acids that do not have non-aromatic hydrocarbon groups (e.g. isophthalic acid, benzoic acid, phthalic acid, 2,5,4'-tricarboxybiphenyl) are unexpectedly harmful components. Although such compounds have significantly reduced chemical activity in the considered reaction medium compared to oxidizable compounds having non-aromatic hydrocarbon groups, it has been found that these compounds, however, undergo many harmful reactions. Therefore, it is useful to control the content of these compounds at preferred intervals in the liquid phase of the reaction medium. This leads to preferred ranges of the selected compounds in the charge of the recycled solvent, as well as to preferred intervals of the selected precursors in the charge of the oxidizable aromatic compound.

For example, in the liquid phase partial oxidation of para-xylene to terephthalic acid (TFA), the applicants found that the intensely colored and undesirable admixture of 2,7-dicarboxyfluorenone (2,7-DCF, 2,7-DCF) is practically not detectable in the reaction medium, and the product is withdrawn when meta-substituted aromatic compounds are present in the reaction medium at very low levels. It is established that when an admixture of isophthalic acid is present in the solvent charge in increased amounts, the formation of 2,7-DCF increases almost in direct proportion. It was also found that when an admixture of meta-xylene is present in the para-xylene charge, the formation of 2,7-DCF also increases almost in direct proportion. In addition, it has been found that even if the solvent charge and the charge of the oxidizable compound are freed from meta-substituted aromatic compounds, some isophthalic acid is formed during the typical partial oxidation of very pure para-xylene, especially when benzoic acid is present in the liquid phase of the reaction Wednesday. Such a self-formed isophthalic acid, due to its higher solubility than TPA, in a solvent consisting of acetic acid and water, can accumulate over time in industrial plants using a recycled solvent. Therefore, the amount of isophthalic acid in the solvent charge, the amount of meta-xylene in the charge of the oxidizable compound and the rate of self-formation of isophthalic acid within the reaction medium are all considered in balance with each other and in balance with any reactions that consume isophthalic acid. Isophthalic acid has been found to undergo other uneconomical reactions, in addition to the formation of 2,7-DCF, which are described below. In addition, it was found that there are other objects to consider when establishing suitable intervals for meta-substituted aromatic samples in the partial oxidation of para-xylene to TFA. Other intensely colored and undesirable impurities, such as 2,6-dicarboxyfluorenone (2,6-DCF), appear to be largely associated with dissolved para-substituted aromatic samples that are always present in the para-xylene charge during liquid phase oxidation. Therefore, suppression of the formation of 2,7-DCF is best considered with regard to the level of other produced colored impurities.

For example, during the liquid-phase partial oxidation of para-xylene to TFA, it was found that the formation of trimellitic acid increases as the levels of isophthalic acid and phthalic acid increase in the reaction medium. Trimellitic acid is a trifunctional carboxylic acid that leads to branching of polymer chains in the production of PET from TFA. In many applications of PET, the degree of branching must be controlled at low levels and, therefore, in purified TFA, trimellitic acid must be controlled at low levels. In addition to the formation of trimellitic acid, the presence of meta-substituted and ortho-substituted samples in the reaction medium also leads to the appearance of other tricarboxylic acids (for example, 1,3,5-tricarboxybenzene). In addition, the increased presence of tricarboxylic acids in the reaction medium increases the formation of tetracarboxylic acids (for example, 1,2,4,5-tetracarboxybenzene). Controlling the total production of all aromatic carboxylic acids having more than two carboxyl groups is one of the factors in establishing the preferred levels of meta-substituted and ortho-substituted samples in the charge of the recycled solvent, in the charge of the oxidizable compound and in the reaction medium in accordance with the present invention .

For example, during the liquid-phase partial oxidation of para-xylene to TFA, it was found that increased levels in the liquid phase of the reaction medium of some dissolved aromatic acids that do not have non-aromatic hydrocarbon groups directly lead to increased production of carbon monoxide and carbon dioxide. Such an increased production of carbon oxides means a loss in the yield of both an oxidizing agent and an oxidizable compound, the latter due to many aromatic carboxylic acids obtained simultaneously, which, on the one hand, can be considered as impurities and, on the other hand, are also of industrial importance . Therefore, the corresponding removal of relatively soluble carboxylic acids without non-aromatic hydrocarbon groups from the recirculated solvent is of economic importance in preventing loss of yield of the oxidizable compound and oxidizing agent, in addition to suppressing the formation of highly undesirable impurities, such as various fluorenones and trimellitic acid.

For example, during the liquid-phase partial oxidation of para-xylene to TFA, it was found that the formation of 2,5,4'-tricarboxybiphenyl seems inevitable. 2,5,4'-Tricarboxybiphenyl is an aromatic tricarboxylic acid formed by combining two aromatic rings, probably due to a combination of dissolved para-substituted aromatic samples with an aryl radical, probably an aryl radical, formed by decarboxylation or decarbonylation of para-substituted aromatic samples. Fortunately, 2,5,4'-tricarboxybiphenyl is usually formed in lower amounts than trimellitic acid, and usually does not lead to significantly increased branching problems of polymer molecules in the production of PET. However, applicants have found that elevated levels of 2,5,4'-tricarboxybibenyl in a reaction medium comprising oxidation of alkyl aromatic compounds in accordance with preferred embodiments of the present invention result in elevated levels of intensely colored and undesired 2,6-DCF. Increased amounts of 2,6-DCF are apparently formed from 2,5,4'-tricarboxybiphenyl due to the opening of the cycle with the loss of a water molecule, although the exact reaction mechanism is not known reliably. If too much 2,5,4'-tricarboxybiphenyl, more soluble in a solvent containing acetic acid and water than TFA, is allowed to accumulate in a recycled solvent, the degree of conversion to 2.6-DCF may become unacceptably large.

For example, during the liquid-phase partial oxidation of para-xylene to TFA, it was found that aromatic carboxylic acids that do not have non-aromatic hydrocarbon groups (e.g., isophthalic acid) usually lead to moderate inhibition of the chemical activity of the reaction medium when they are present in the liquid phase in a sufficient concentration.

For example, during liquid-phase partial oxidation of para-xylene to TFA, it was found that deposition is very often imperfect (i.e., nonequilibrium) with respect to the corresponding concentrations of various chemical samples in the solid phase and in the liquid phase. This is probably due to the fact that the deposition rate is very high at bulk reaction rates, preferred in this case, which leads to imperfect coprecipitation of impurities or even to occlusion. Therefore, when it is desirable to limit the concentration of certain impurities (for example, trimellitic acid and 2,6-DCF) in crude TFA, due to the configuration of the downstream installation, it is preferable to control their concentration in the solvent charge, as well as the rate of their formation in the reaction medium .

For example, it was found that benzophenone compounds (e.g., 4,4'-dicarboxybenzophenone and 2,5,4'-tricarboxybenzophenone) obtained during the partial oxidation of para-xylene have an undesirable effect on the reaction medium of PET, even though the benzophenone compounds in TFA are not in themselves intensely colored, like fluorenones and anthraquinones. Accordingly, it is desirable to limit the presence of benzophenones and selected precursors in the charge of the recycled solvent and in the charge of the oxidizable compound. In addition, it was found that the presence of elevated levels of benzoic acid, regardless of whether it is in a recycled solvent or formed in the reaction medium, leads to increased production rates of 4,4'-dicarboxybenzophenone.

In the course of the study, the applicants established and quantified an unexpected reaction order for aromatic compounds that do not have non-aromatic hydrocarbon groups that are present during the liquid-phase partial oxidation of para-xylene to TFA. When analyzing only one case of benzoic acid, the applicants found that increased levels of benzoic acid in the reaction medium of some embodiments of the present invention lead to significantly increased production of intensely colored and undesired 9-fluorenone-2-carboxylic acid, to significantly increased levels of 4,4'- dicarboxybiphenyl, to elevated levels of 4,4'-dicarboxybenzophenone, to moderate inhibition of chemical activity in the direction of the target oxidation of para-xylene, as well as to elevated levels carbon dioxide and associated yield loss. It has been found that elevated levels of benzoic acid in the reaction medium also lead to increased production of isophthalic acid and phthalic acid, the levels of which it is desirable to control at low intervals in accordance with similar aspects of the present invention. The number and significance of reactions involving benzoic acid is probably even more unexpected, as some modern researchers consider the use of benzoic acid instead of acetic acid as the main component of the solvent (see, for example, US patent No. 6562997). In addition, it is observed that benzoic acid self-forms during the oxidation of para-xylene at rates that are sufficiently significant relative to its formation from impurities such as toluene and ethylbenzene, usually found in the loading of an oxidizable compound containing industrially pure para-xylene.

On the other hand, applicants have disclosed a small value for further adjusting the composition of the recycled solvent with respect to the presence of an oxidizable aromatic compound and with respect to the reactions of aromatic intermediates that contain non-aromatic hydrocarbon groups and are relatively soluble in the recycled solvent. In the General case, such compounds are either fed into the reaction medium, or formed in the reaction medium at significantly higher speeds than their amount present in the recycled solvent; and the rate of consumption of these compounds in the reaction medium is high enough to hold one or more non-aromatic hydrocarbon groups to limit their accumulation in the recycled solvent. For example, during partial oxidation of para-xylene in a multiphase medium, para-xylene evaporates to a limited extent along with large amounts of solvent. When such an evaporated solvent leaves the reactor as part of the effluent gas and condenses to be recovered as a recycled solvent, a significant part of the vaporized para-xylene also condenses in it. There is no need to limit the concentration of such para-xylene in the recycled solvent. For example, if the solvent is separated from solids when the suspension leaves the para-xylene oxidation reaction medium, this isolated solvent will contain a concentration of dissolved para-toluic acid similar to that at the point of removal from the reaction medium. Although it may be important to limit the fixed concentration of para-toluic acid in the liquid phase of the reaction medium (see below), it is not necessary to separately regulate para-toluic acid in this portion of the recycled solvent due to its relatively good solubility and its low specific mass flow rate relative to the formation of para- toluic acid within the reaction medium. Similarly, insignificant reasons have been identified for limiting the concentration in a recycled solvent of aromatic compounds with methyl substituents (e.g. toluyl acids), aromatic aldehydes (e.g. terephthalic aldehyde), aromatic compounds with hydroxymethyl substituents (e.g. 4-hydroxymethylbenzoic acid) and brominated aromatic compounds containing at least one non-aromatic hydrocarbon group (e.g. alpha bromo-para-toluic acid) is lower than the concentration In fact found in the liquid phase exiting from the reaction medium, existing in the partial oxidation of xylene according to preferred embodiments of the present invention. Unexpectedly, the applicants found that there was also no need to control the concentration of selected phenols in the recycled solvent, which were essentially produced during the partial oxidation of xylene, since these compounds formed and decomposed within the reaction medium at rates much higher than their presence in the recycled solvent. For example, it has been found that 4-hydroxybenzoic acid has a relatively small effect on the reactivity in preferred embodiments of the present invention when co-fed at a rate of over 2 grams of 4-hydroxybenzoic acid per 1 kilogram of para-xylene, significantly higher than the natural presence in the recycled solvent despite the fact that other authors point to it as a strong poison in a similar reaction medium (see, for example, W. Partenheimer, Catalysis Today 23, (1995), p. 81).

Thus, there are many reactions and numerous considerations in establishing the preferred ranges of various aromatic impurities in the solvent charge, as currently shown. Such results are presented in terms of the total weighted average composition of all solvent streams that are supplied to the reaction medium for a predetermined period of time, preferably within one day, more preferably one hour and most preferably within 1 minute. For example, if one solvent charge flows substantially continuously with a composition of 40 parts by weight of isophthalic acid at a rate of 7 kg / min, a second solvent charge flows substantially continuously with a composition of 2000 parts by weight of isophthalic acid at a rate of 10 kg / min, and there are no other solvent loading streams entering the reaction medium, the total weighted average solvent loading composition is calculated as (40 * 7 + 2000 * 10) / (7 + 10) = 1.193 mass.ppm of isophthalic acid. It should be noted that the loading mass of any oxidizable compound or the loading of any oxidizing agent, which may be combined with the loading of the solvent before entering the reaction medium, are not taken into account when calculating the total weighted average solvent loading composition.

Table 1 below provides preferred values for some of the components in the loading of solvent introduced into the reaction medium. The solvent loading components shown in Table 1 are the following compounds: 4-carboxybenzaldehyde (4-KBA), 4,4'-dicarboxylstilbene (4,4'-DCS, 4,4'-DCS), 2,6-dicarboxyanthraquinone (2,6-DKA, 2,6-DCA), 2,6-dicarboxyfluorenone (2,6-DKF), 2,7-dicarboxyfluorenone (2,7-DKF), 3,5-dicarboxyfluorenone (3,5- DCF), 9-fluorenone-2-carboxylic acid (9F-2KK, 9F-2CA), 9-fluorenone-4-carboxylic acid (9F-4KK, 9F-4CA), all fluorenones, including other fluorenones not listed separately (all fluorenones), 4,4'-dicarboxybiphenyl (4,4'-DCB), 2,5,4'-tricarboxybiphenyl (2,5,4'-TKB, 2,5,4'-TCB), ft Aleic acid (FC, RA), isophthalic acid (IPA, IPA), benzoic acid (BC, VA), trimellitic acid (TMK, TMA), 2,6-dicarboxybenzocoumarin (2,6-DCBC, 2,6-DCBC) 4,4'-dicarboxybenzyl (4,4'-DCBZ, 4,4'-DCBZ), 4,4'-dicarboxybenzophenone (4,4'-DCBF, 4,4'-DCBP), 2,5,4 '-tricarboxybenzophenone (2,5,4'-TKBF, 2,5,4'-ТСВР), terephthalic acid (TFA), solids precipitated at 20 ° С and all aromatic carboxylic acids without non-aromatic hydrocarbon groups. The following table 1 shows the preferred amounts of these impurities in STK produced in accordance with an embodiment of the present invention.

TABLE 1
The components of the solvent loading introduced into the reaction mixture

Figure 00000004

Figure 00000005

Many other aromatic impurities, as a rule, are also present in the recycled solvent, and usually their quantities vary even at lower levels and / or in proportion to one or more of the disclosed aromatic compounds. Methods for controlling the disclosed aromatic compounds at preferred ranges will typically keep other aromatic impurities at suitable levels.

When bromine is used in a reaction medium, large quantities of ionic and organic forms of bromine are known to be in dynamic equilibrium. These various forms of bromine have different stability characteristics when exiting the reaction medium and when passing through various processing units related to the recycled solvent. For example, alpha-bromo-para-toluic acid may continue to exist as such under certain conditions, or may rapidly hydrolyze under other conditions to form 4-hydroxymethylbenzoic acid and hydrogen bromide. In the present invention, it is preferable that at least about 40% of the mass., More preferably, at least about 60% of the mass. and most preferably, at least about 80% of the mass. The total mass of bromine present in the total solvent charge in the reaction medium was in one or more of the following chemical forms: ionic bromine, alpha-bromo-para-toluic acid, and bromoacetic acid.

Although the importance and importance of controlling the total weighted average solvent loading purity within the disclosed, desired ranges of the present invention have not yet been discovered and / or disclosed, suitable solvent purity control means can be collected from various methods already known in the art. First, any solvent vaporized from the reaction medium is generally of a suitable purity, provided that the liquid or solids from the reaction medium are not captured by the vaporized solvent. The supply of reflux from droplets of solvent into the separation space of the outgoing gas above the reaction medium, as described in this case, appropriately limits such capture; and from this effluent gas, a recycled solvent of suitable purity with respect to the aromatic compound may be condensed. Secondly, a more difficult and expensive cleaning of the charge of the recirculated solvent is associated with a solvent selected from the reaction medium in liquid form and with a solvent which subsequently comes into contact with the liquid and / or solid phases of the reaction medium withdrawn from the reaction vessel (e.g., recycled solvent obtained from a filter on which solids were concentrated and / or washed, a recycled solvent obtained from a centrifuge on which solids were concentrated and / or washed recycled solvent selected from the crystallization operation, etc.). However, means are also well known in the art for efficiently performing the necessary purification of such recycled solvent streams using one or more of the prior art discoveries. Regarding the control of precipitated solids in a recycled solvent within specially defined intervals, suitable control tools include, but are not limited to, gravimetric sedimentation, mechanical filtration using filter cloth on rotary belt filters and rotary drum filters, mechanical filtration using a stationary filter media in pressure vessels, hydrocyclones and centrifuges. Regarding the control of dissolved aromatic samples in a recycled solvent within specially defined ranges, control agents include, but are not limited to, the agents described in US Pat. No. 4,939,297 and US Patent Application Publication Number 2005-0038288, which are incorporated by reference. . However, none of the preceding inventions discloses or describes the preferred purity levels of the total solvent charge disclosed herein. Rather, these prior art inventions offer only means for purifying selected and partial recycled solvent streams without disclosing the present invention, the optimal composition values of the total weighted average loading of the solvent into the reaction medium.

Returning to the purity of the charge of the oxidizable compound, it is known that certain levels of isophthalic acid, phthalic acid and benzoic acid are present and at low concentrations are acceptable in the purified TFA used to make the polymers. Moreover, it is known that these samples have a relatively higher solubility in many solvents and can be successfully removed from purified TFA using crystallization processes. However, from the embodiment of the invention described in this invention, it is known that controlling the level of some relatively soluble aromatic samples, especially including isophthalic acid, phthalic acid and benzoic acid, in the liquid phase of the reaction medium is unexpectedly of great importance for controlling the level of polycyclic and colored aromatic compounds formed in the reaction medium to control compounds with more than 2 carboxylic acid functions in the molecule, to control the reaction activity within the partial oxidation reaction medium and to control the yield loss of the oxidizing agent and aromatic compound.

It is known in the art that isophthalic acid, phthalic acid and benzoic acid are formed in the reaction medium as follows. Impurities of meta-xylene in the charge are oxidized with good conversion and give IFC. Impurities of ortho-xylene in the charge are oxidized with good conversion and give phthalic acid. Impurities of ethylbenzene and toluene in the charge are oxidized with good conversion and give benzoic acid. However, significant amounts of isophthalic acid, phthalic acid and benzoic acid have also been found to be formed within the reaction medium containing para-xylene by methods other than the oxidation of meta-xylene, ortho-xylene, ethylbenzene and toluene. These other internal chemical pathways probably include decarbonylation, decarboxylation, rearrangement of transition states, and the addition of methyl and carbonyl radicals to aromatic rings.

In determining the preferred ranges of impurities in the loading of the oxidizable compound, many factors are significant. Any impurity in the charge appears to result in a direct loss of yield and in product cleaning costs if the purity requirements for the oxidized product are quite stringent (for example, in the reaction medium for partial oxidation of para-xylene, toluene and ethylbenzene are usually found in para- xylene of industrial purity, give benzoic acid, and this benzoic acid is mainly removed from most of the industrial TFA). When the product of partial oxidation of charge impurities precipitates during additional reactions, factors other than simple loss of yield and removal become significant when considering what increase in the cost of purification of raw materials results from this (for example, in the reaction medium of partial oxidation of para-xylene, ethylbenzene leads to benzoic acid, and benzoic acid then gives intensely colored 9-fluorenone-2-carboxylic acid, isophthalic acid, phthalic acid and an increased amount of carbon oxides, among others mi). When the reaction medium itself generates additional amounts of impurities due to chemical mechanisms not directly related to the impurities of the charge, the analysis becomes even more complicated (for example, in the reaction medium for the partial oxidation of para-xylene, benzoic acid also forms from para-xylene itself). In addition, downstream processing of the crude oxidation product may affect consideration of the preferred purity of the feed. For example, the cost of removal to acceptable levels of direct impurity (benzoic acid) and the cost of removing subsequent impurities (isophthalic acid, phthalic acid, 9-fluorenone-2-carboxylic acid, etc.) may be the same, may differ from each other, and may differ from the requirements for the removal of most unrelated impurities (for example, the product of incomplete oxidation of 4-CBA during the oxidation of para-xylene to TFA).

The purity loading ranges disclosed below for para-xylene are preferred, where para-xylene is fed with a solvent and an oxidizing agent to the partial oxidation reaction medium to produce TFA. These ranges are more preferable in the TFA manufacturing process, including subsequent oxidation steps to remove impurities other than the oxidizing agent and solvent (e.g., catalyst metals) from the reaction medium. Such ranges still remain preferable in TFA production processes in which additional 4-CBA is removed from the STC (for example, by converting the STC to dimethyl terephthalate plus impurity esters and subsequent removal of the 4-CBA methyl ester by distillation using oxidative combustion processes to convert 4-CBA in TFA, using hydrogenation methods to convert 4-CBA to para-toluic acid, which is then separated by fractional crystallization). These ranges are most preferred in TFA production processes in which additional 4-KBA is removed from the STK using oxidative combustion processes to convert 4-KBA to TFA.

Using new knowledge of the preferred ranges of recycle aromatic compounds and the relative amounts of aromatic compounds formed directly during the oxidation of the charge impurities, compared with other unrelated chemical pathways, improved impurity ranges for contaminated para-xylene, which are fed to the partial oxidation process to produce TFK. Table 2 below presents the preferred values for the amount of meta-xylene, ortho-xylene and ethylbenzene + toluene in the para-xylene charge, expressed in parts by weight per million para-xylene.

TABLE 2
Contaminated Para-Xylene Charge Components
Component Identification Preferred Amounts (ppm) More preferred amounts (parts by weight per million) Most preferred amounts (parts by weight per million) Meta-Xylene 20-800 50-600 100-400 Ortho-xylene 10-300 20-200 30-100 Ethylbenzene + toluene * 20-700 50-500 100-300 Total 50-900 100-800 200-700 * Definition for (ethylbenzene + toluene) - each separately and in total

A person skilled in the art will now understand that the above impurities in contaminated para-xylene can have the strongest effect on the reaction medium after the accumulation of products of their partial oxidation in a recycled solvent. For example, supplying an upper amount from the most preferred meta-xylene range, 400 ppmw, will immediately produce about 200 ppm isophthalic acid in the liquid phase of the reaction medium when operating with about 33 wt%. solids in the reaction medium. This is comparable to entering from the upper amount of the most preferred interval for isophthalic acid in a recycled solvent, 400 parts by mass per million, which, after carrying out a typical evaporation of the solvent to cool the reaction medium, reaches approximately 1200 parts by mass of isophthalic acid in the liquid phase reaction medium. Therefore, it is the accumulation of partial oxidation products over time in a recycled solvent that is the most likely effect of impurities of meta-xylene, ortho-xylene, ethylbenzene and toluene in the loading of contaminated para-xylene. Accordingly, the above impurity intervals in the supply of contaminated para-xylene are preferably maintained for at least half a day every day of any partial oxidation reaction medium in a particular industrial plant, more preferably at least three quarters of a day every day for at least at least seven consecutive work days and most preferably when the weight-average contaminated para-xylene loading composition is within the preferred ranges of at least m Here, for 30 consecutive work days.

Means for producing contaminated para-xylene of preferred purity are already known in the art and include, but are not limited to, distillation, fractional crystallization methods at temperatures below ambient temperature and molecular sieve methods using pore-selective adsorption. However, the preferred purity ranges defined in the present invention at their high boundary are more costly and expensive than is usually the case with industrial feeds of para-xylene; and furthermore, at a low boundary, the preferred ranges exclude the overly expensive purification of para-xylene to provide partial oxidation to the medium through detection and disclosure, where the combined effect of the self-formation of impurities from para-xylene itself and the impurity-eliminating reactions within the reaction medium become more more important than the feed rates of impurities in contaminated para-xylene.

When the xylene-containing feed stream contains selected impurities, such as ethylbenzene and / or toluene, the oxidation of such impurities can produce benzoic acid. As used herein, the term “benzoic acid formed from impurities” will mean benzoic acid obtained from any source other than xylene during the oxidation of xylene.

As stated in this invention, a portion of the benzoic acid generated during the oxidation of xylene is derived from xylene itself. Such a production of benzoic acid from xylene is obviously added to any part of the production of benzoic acid, which may be formed from impurities with benzoic acid. Without being attached to any theory, it is believed that benzoic acid is obtained from xylene in a reaction medium when various products of intermediate xylene oxidation spontaneously undergo decarbonylation (loss of carbon monoxide) or decarboxylation (loss of carbon dioxide), resulting in the formation of aryl radicals. Such aryl radicals can then subtract a hydrogen atom from one of the many available sources in the reaction medium and produce self-generated benzoic acid. Regardless of the chemical mechanism, the definition of "self-formed benzoic acid" used in this case will mean benzoic acid obtained from xylene during the oxidation of xylene.

As discussed above, when para-xylene is oxidized to form terephthalic acid (TFA), the production of self-formed benzoic acid leads to a loss of para-xylene yield and a loss of oxidative yield. In addition, the presence of self-formed benzoic acid in the liquid phase of the reaction medium is consistent with the growth of many undesirable side reactions, in particular including the formation of intensely colored compounds called monocarboxyfluorenones. Self-formed benzoic acid also contributes to the undesired accumulation of benzoic acid in the recycled solvent, which further increases the concentration of benzoic acid in the liquid phase of the reaction medium. Consequently, the production of self-formed benzoic acid is desirable to be minimized, but it is also appropriately taken into account simultaneously with benzoic acid formed from impurities, with factors affecting the consumption of benzoic acid, with factors related to other issues of reaction selectivity, and with the overall economics of the process.

It has been found that self-generated benzoic acid can be controlled at low levels by appropriate selection, for example, temperature, xylene distribution and oxygen availability within the reaction medium during oxidation. Without being bound by any theory, lower temperatures and improved oxygen availability appear to suppress decarbonylation and / or decarboxylation rates, thus eliminating the aspect of yield loss due to self-formed benzoic acid. Sufficient accessibility, it turns out, directs aryl radicals to the formation of other less dangerous products, in particular hydroxybenzoic acids. The distribution of xylene in the reaction medium can also affect the balance between the conversion of the aryl radical to benzoic acid or to hydroxybenzoic acids. Regardless of the chemical mechanisms, applicants have disclosed reaction conditions that, although gentle enough to reduce the production of benzoic acid, are harsh enough to oxidize a high proportion of the produced hydroxybenzoic acids to carbon monoxide and / or carbon dioxide, which are easily removed from the oxidation product.

In a preferred embodiment of the present invention, the oxidation reactor is configured and operates in such a way that the formation of self-generated benzoic acid is minimal and the oxidation of hydroxybenzoic acids to carbon monoxide and / or carbon dioxide is minimized. When the oxidation reactor is used to oxidize para-xylene to terephthalic acid, it is preferred that para-xylene is at least about 50% by weight. from all xylenes in the feed stream introduced into the reactor. More preferably, para-xylene is at least about 75% of the mass. from all xylene in the feed stream. Even more preferably, para-xylene is at least about 95% of the mass. from all xylene in the feed stream. Most preferably, para-xylene comprises substantially all of the xylene in the feed stream.

When the reactor is used to oxidize para-xylene to terephthalic acid, it is preferable that the production rate of terephthalic acid is maximum, while the production rate of self-formed benzoic acid is minimal. Preferably, the ratio of the production rate (by weight) of terephthalic acid to the production rate (by weight) of self-formed benzoic acid is at least about 500: 1, more preferably at least about 1000: 1, and most preferably at least , 1500: 1. As will be shown below, the production rate of the self-formed benzoic acid is preferably measured when the concentration of benzoic acid in the liquid phase of the reaction medium is below 2000 parts by mass / million, more preferably below 1000 parts by mass / million and most preferably below 500 parts by mass. / million, since such low concentrations are suppressed to reasonably low reaction rates at which benzoic acid is converted to other compounds.

When combining self-generated benzoic acid and benzoic acid formed from impurities, the ratio of the production rate (by weight) of terephthalic acid to the production rate (by weight) of all (self-formed and formed from impurities) benzoic acid is preferably at least about 400: 1, more preferably at least about 700: 1 and most preferably at least 1100: 1. As will be shown below, the summed production rate of self-formed benzoic acid plus formed from impurities of benzoic acid is preferably measured when the concentration of benzoic acid in the liquid phase of the reaction medium is lower than 500 ppm, since such low concentrations are suppressed to reasonably low reaction rates by which benzoic acid is converted into other compounds.

As shown in this paper, increased concentrations of benzoic acid in the liquid phase of the reaction medium lead to increased formation of many other aromatic compounds, some of which are harmful impurities in TFA; and, as shown in the present invention, increased concentrations of benzoic acid in the liquid phase of the reaction medium lead to increased formation of gaseous carbon oxides, the formation of which means loss of yield in the oxidizing agent and in the aromatic compounds and / or solvent. Moreover, it has now been established that a significant part of this increased formation of other aromatic compounds and carbon oxides originates from reactions in which some molecules of benzoic acid itself are converted, in contrast to benzoic acid, which catalyzes other reactions without its own consumption. Accordingly, “pure benzoic acid formation” is defined in this case as the time-averaged mass of all benzoic acid leaving the reaction medium, minus the time-averaged mass of all benzoic acid entering the reaction medium during the same time period. Such a pure formation of benzoic acid is often positive, resulting from the rates of formation of benzoic acid formed from impurities and self-formed benzoic acid. However, it was found that the rate of conversion of benzoic acid to carbon oxides and to some other compounds, it turns out, increases approximately linearly with increasing concentration of benzoic acid in the liquid phase of the reaction medium, measured when other reaction conditions, including temperature, oxygen availability, SOS and reactivity are suitably maintained constant. Therefore, when the concentration of benzoic acid in the liquid phase of the reaction medium is sufficiently high, probably due to the increased concentration of benzoic acid in the recycled solvent, the conversion of benzoic acid molecules to other compounds, including carbon oxides, may become equal to or greater than the chemical the formation of new benzoic acid molecules. In this case, the pure formation of benzoic acid can balance near zero or even be negative. It has been found that when the net formation of benzoic acid is positive, the ratio of the production rate (by weight) of terephthalic acid in the reaction medium to the rate of pure formation of benzoic acid in the reaction medium is preferably more than about 700: 1, more preferably about more than 1100: 1 and most preferably about 4000: 1. Applicants have found that when the net formation of benzoic acid is negative, the ratio of the production rate (by weight) of terephthalic acid in the reaction medium to the rate of pure formation of benzoic acid in the reaction medium is preferably more than about 200: (- 1), more preferably about more than 1000: (- 1) and most preferably about 5000: (- 1).

Applicants also disclose preferred ranges for suspension composition (liquid + solids) discharged from the reaction medium, and for a portion of the solid STK in suspension. Compositions of the preferred suspension and preferred STK are unexpectedly broader and more useful. For example, purified TFA produced from such a preferred STC by oxidative combustion has a fairly low content of all impurities and colored impurities, therefore, purified TFA is acceptable without hydrogenation of an additional amount of 4-KBA and / or colored impurities for a wide range of applications in PET fibers and PET packaging . For example, the composition of the preferred suspension gives a liquid phase of the reaction medium, which has a relatively low concentration of important impurities, and this significantly reduces the formation of other even more undesirable impurities, as described in the present invention. In addition, the composition of the preferred suspension greatly facilitates the subsequent processing of the liquid from the suspension so that it becomes an acceptable pure recycled solvent in accordance with other embodiments of the present invention.

An STK produced in accordance with one embodiment of the present invention contains fewer impurities of established types than an STK produced using conventional processes and equipment, especially those using a recycled solvent. Impurities that may be present in the STC include the following compounds: 4-carboxybenzaldehyde (4-CBA), 4,4'-dicarboxylstilbene (4,4'-DCS, 4,4'-DCS), 2,6-dicarboxyanthraquinone (2 6-DKA, 2,6-DCA), 2,6-dicarboxyfluorenone (2,6-DKF), 2,7-dicarboxyfluorenone (2,7-DKF), 3,5-dicarboxyfluorenone (3,5-DKF) 9-fluorenone-2-carboxylic acid (9F-2KK, 9F-2CA), 9-fluorenone-4-carboxylic acid (9F-4KK, 9F-4CA), all fluorenones, including other fluorenones not listed separately (all fluorenones ), 4,4'-dicarboxybiphenyl (4,4'-DCB), 2,5,4'-tricarboxybiphenyl (2,5,4'-TKB, 2,5,4'-TSB), phthalic acid (FC, RA) isophthalic acid (IPA, IPA), benzoic acid (BC, VA), trimellitic acid (TMK, TMA), para-toluic acid (PTK, PTAC), 2,6-dicarboxybenzocoumarin (2,6-DCBA, 2,6- DCBC), 4,4'-dicarboxybenzyl (4,4'-DCBZ, 4,4'-DCBZ), 4,4'-dicarboxybenzophenone (4,4'-DCBF, 4,4'-DCBP), 2.5 , 4'-tricarboxybenzophenone (2,5,4'-TKBF, 2,5,4'-ТСВР). The following table 3 shows the preferred amounts of these impurities in STK produced in accordance with an embodiment of the present invention.

TABLE 3
Impurities in STK

Figure 00000006

In addition, it is preferable that the STK produced in accordance with an embodiment of the present invention has a lower content of colored components compared to the STK produced using conventional processes and equipment, especially those using a recycled solvent. Therefore, it is preferable that the STK produced in accordance with one embodiment of the present invention have a transmittance at 340 nanometers (nm) of at least about 25%, more preferably at least about 50%, and most preferably at least 60%. It is also preferred that the CTK produced in accordance with one embodiment of the present invention has a transmittance at 400 nm of at least about 88%, more preferably at least about 90%, and most preferably at least 92%.

The percent transmission test provides a measure of the amount of colored, light-absorbing impurities present in TFA or STK. The test used in this case refers to measurements made on a portion of a solution prepared by dissolving 2.00 g of dry solid TFA or STK in 20.0 ml of analytical dimethyl sulfoxide (DMSO) or higher. A portion of this solution is placed in a Hellma semi-micro-flow cell, PN 176.700, which is made of quartz and has a path length of 1.0 cm and a volume of 0.39 ml (Hellma USA, 80 Skyline Drive, Plainview, NY 11803). An Agilent 8453 Diode Array Spectrophotometer is used to measure the transmission of various wavelengths of light through a filled flow cell (Agilent Technologies, 395 Page Mill Road, Palo Alto, CA 94303). After appropriate correction of the absorption relative to the background, including, but not limited to, the cell and the solvent used, the percent transmission data characterizing the fraction of the incident light that passes through the solution is recorded directly using the device. The transmission percentages at wavelengths of 340 and 400 nm are especially useful for separating pure TFA from many of the impurities commonly found in it.

Preferred ranges of various aromatic impurities in the suspension phase (solids + liquid) of the reaction medium are presented below in table 4.

TABLE 4
Impurities in suspension

Figure 00000007

Such preferred suspension compositions mean a preferred liquid phase composition of the reaction medium, while avoiding experimental problems associated with the deposition of additional liquid phase components from the reaction medium into solid phase components during sampling from the reaction medium, separation of liquids and solids, and bias analytical conditions.

Many other aromatic impurities, as a rule, are also present in the suspension phase of the reaction medium and in the STK of the reaction medium, and their amount usually varies at lower levels and / or in proportion to one or more disclosed aromatic compounds. Controlling the disclosed aromatic compounds at preferred ranges will maintain other aromatic impurities at acceptable levels. Such preferred compositions of the suspension phase in the reaction medium and in solid STK taken directly from the suspension make it possible to work with the embodiments of the present invention described for the partial oxidation of para-xylene to TFA.

Measurement of the concentration of components present at a low level in a solvent, a recycled solvent, STK, suspensions from the reaction medium and TFA are carried out using liquid chromatography methods. Two interchangeable embodiments are described below.

A method called HPLC-DAD involves high performance liquid chromatography (HPLC) in combination with diode array detection to allow separation and quantification of various types of molecules in a given sample. The instrument used for such measurements is an 1100 HPLC equipped with a DAD (diode array detector) available from Agilent Technologies (Palo Alto, CA), although other suitable instruments are also commercially available from other suppliers. As is known in the art, both the elution time and the response of the detector are calibrated using known compounds present in known amounts, compounds and amounts that correspond to compounds and quantities found in unknown samples.

A method called HPLC-MS includes high performance liquid chromatography (HPLC) in combination with mass spectrometry (MS) to allow separation, identification and quantification of various types of molecules in a given sample. The instrument used for such measurements is the Alliance HPLC and ZQ MS offered by Waters Corp. (Milford, MA), although other suitable appliances are also commercially available from other suppliers. As is known in the art, both the elution time and the mass spectrometric response are calibrated using known compounds present in known amounts, compounds and amounts that correspond to compounds and amounts found in unknown samples.

Another embodiment of the present invention relates to the partial oxidation of an aromatic, oxidizable compound with appropriate balancing of the suppression of harmful aromatic impurities, on the one hand, with respect to the production of carbon monoxide and carbon dioxide (collectively carbon oxides (CO x )), on the other hand. Such carbon oxides typically exit the reaction vessel as off-gas, and they mean a destructive loss of solvent and oxidizable compound, including, of course, and, ultimately, preferred oxidized derivatives (e.g., acetic acid, steam -xylene and TPA). Applicants have identified lower limits for the formation of carbon oxides, below which, as it turns out, the high formation of harmful impurities, which are described below, and the level of low conversion are inevitably too bad for economic applicability. The upper boundaries of carbon dioxide are also disclosed, above which the formation of carbon dioxide continues to grow with a small additional value provided by a decrease in the formation of harmful aromatic impurities.

Applicants have also found that lowering liquid-phase concentrations of the charge of the oxidizable aromatic compound and aromatic intermediate samples in the reaction medium leads to lower formation rates of harmful impurities during the partial oxidation reaction of the oxidizable aromatic compound. Such harmful impurities include conjugated aromatic rings and / or aromatic molecules containing more than the desired number of carboxylic acid groups (for example, when para-xylene is oxidized, 2,6-dicarboxyananthraquinone, 2,6-dicarboxyfluorenone, trimellitic acid, 2,5 are harmful , 4'-tricarboxybiphenyl and 2,5,4'-benzophenone). Aromatic intermediates include aromatic compounds that have passed from the charge of an oxidizable compound and still retain non-aromatic hydrocarbon groups (for example, when para-xylene is oxidized, aromatic intermediates are para-toluyl aldehyde, terephthalic aldehyde, para-toluic acid, 4-CBA, 4-hydroxymethylbenzoic acid and alpha-bromo-para-toluic acid). The loading of an oxidizable aromatic compound and aromatic intermediate samples retaining non-aromatic hydrocarbon groups when they are present in the liquid phase of the reaction medium, as it turns out, lead to harmful impurities in a manner similar to that already described for dissolved aromatic samples without non-aromatic hydrocarbon groups ( e.g. isophthalic acid).

Contrasting such a need for higher reactivity to suppress the formation of harmful aromatic impurities during the partial oxidation of oxidizable aromatic compounds, the applicants have found that an undesirable concomitant result is increased production of carbon oxides. It is important to note that such carbon oxides mean a loss in yield of the oxidizable compound and oxidizing agent, and not just the solvent. It is clear that a significant and sometimes fundamental fraction of carbon oxides comes from an oxidizable compound and its derivatives, and not from a solvent, and often an oxidizable compound is more expensive per unit carbon than a solvent. In addition, it is important to note that the desired carboxylic acid product (for example, TFA) also undergoes excessive oxidation to carbon oxides when present in the liquid phase of the reaction medium.

It is also important to note that the present invention relates to reactions in the liquid phase of the reaction medium and to the concentrations of the reactants therein. This contradicts some of the inventions of the prior art, which relate directly to the formation of precipitated solids from aromatic compounds that retain non-aromatic hydrocarbon groups. Specifically, in the case of partial oxidation of para-xylene to TFA, some prior art inventions relate to the amount of 4-CBA precipitated in the solid phase of STC. However, the applicants of the present invention disclose a change from more than two to one for the ratio of 4-KBA in the solid phase to 4-KBA in the liquid phase using a list of temperature, pressure, catalyst, solvent composition and volumetric reaction rate of para-xylene, depending on whether partial oxidation is carried out in a well-stirred autoclave or in a reaction medium with a stepwise change in the concentrations of oxygen and para-xylene in accordance with the present invention. In addition, the applicants observed that the ratio of 4-CBA in the solid phase to 4-CBA in the liquid phase can also vary from two to one or in a well-mixed or stepwise reaction medium, depending on the volumetric rate of the para-xylene reaction with another a similar list of temperature, pressure, catalyst and solvent composition. In addition, 4-KBA in the solid phase of STK, it turns out, does not contribute to the formation of harmful impurities, and 4-KBA in the solid phase can be isolated and oxidized to TFA and in high yield (for example, by oxidative burning of a suspension of STK, as described In that work); while the removal of harmful impurities is much more difficult and expensive than the removal of solid-phase 4-KBA, and the formation of carbon oxides means a constant loss of yield. Thus, it is important to distinguish that this aspect of the present invention relates to liquid phase compositions in a reaction medium.

Regardless of whether the source is a solvent or an oxidizable compound, applicants have found that for conversions acceptable in industrial applications, the production of carbon oxides is closely related to the level of overall reactivity, despite a wide variation in a certain combination of metal temperature, halogen, temperature , the acidity of the reaction medium, measured by pH, the concentration of water used to obtain a certain level of overall reaction activity. It has been found that in the case of partial xylene oxidation, it is useful to increase the level of the general reaction activity using the concentration of toluic acid in the liquid phase at the average height of the reaction medium, at the bottom of the reaction medium and at the top of the reaction medium.

Thus, the importance of simultaneous alignment follows from this in order to minimize the formation of harmful impurities by increasing the reactivity and to also minimize the formation of carbon oxides by reducing the reactivity. That is, if the total production of carbon oxides is suppressed too little, then excessive levels of harmful impurities are formed, and vice versa.

In addition, the applicants have found that the solubility and relative reactivity of the target carboxylic acid (e.g., TFA) and the presence of other dissolved aromatic samples that do not have non-aromatic hydrocarbon groups provide a very important fulcrum for this alignment of carbon oxides with respect to harmful impurities. The desired product is a carboxylic acid, as a rule, dissolved in the liquid phase of the reaction medium, even when present in solid form. For example, at temperatures in the preferred range, TFA dissolves in a reaction medium containing acetic acid and water in amounts ranging from about one thousandth ppm to over 1% by weight, and solubility increases with temperature. Although there are differences in reaction rates in the direction of formation of various harmful impurities from the charge of an oxidizable aromatic compound (e.g., para-xylene), from aromatic intermediates (e.g., para-toluic acid), from the target aromatic carboxylic acid product (e.g., TFA ) and from aromatic samples that do not have non-aromatic hydrocarbon groups (for example, isophthalic acid), the presence and reactivity of the last two groups determine the region of diminishing responses in a relatively further suppression of the first two groups capable of oxidizing the aromatic compound and aromatic intermediate reaction compounds. For example, in the case of partial oxidation of para-xylene to TPA, if the dissolved TPA reaches up to 7000 parts per million in the liquid phase of the reaction medium under these conditions, the dissolved benzoic acid reaches up to 8000 parts per million, and the dissolved isophthalic acid reaches up to 6000 parts by mass per million and dissolved phthalic acid reaches up to 2000 parts by mass per million, then the importance of reducing the level of all harmful compounds begins to decrease as the reaction activity increases to suppress the liquid-phase concentration of para-toluic acid and 4-KBA below similar levels . That is, the presence and concentration in the liquid phase of the reaction medium of aromatic samples without aromatic hydrocarbon groups varies very slightly due to an increase in reaction activity, and their presence serves to expand upward the region of diminishing responses to reduce the concentration of reaction intermediates in order to suppress the formation of harmful impurities.

Therefore, one embodiment of the present invention provides preferred ranges of carbon oxides (carbon monoxide and carbon dioxide), limited at the lower end by low reactivity and excessive formation of harmful impurities, and at the upper end by excessive carbon losses, but at levels lower than previously disclosed and described as useful in industrial implementation. Accordingly, the formation of carbon oxides is preferably controlled as follows. The ratio of moles of all carbon monoxide produced to moles of the oxidizable aromatic compound supplied is preferably in the range of about 0.02: 1 to 0.25: 1, more preferably in the range of about 0.04: 1 to 0.22: 1, even more preferably in the range of from about 0.05: 1 to 0.19: 1, and most preferably in the range of from 0.06: 1 to 0.15: 1. The ratio of moles of carbon dioxide formed to moles of the oxidizable aromatic compound supplied is preferably in the range of about 0.01: 1 to 0.21: 1, more preferably in the range of about 0.03: 1 to 0.19: 1, more more preferably in the range of about 0.04: 1 to 0.16: 1; and most preferably in the range of 0.05: 1 to 0.11: 1. The ratio of moles of carbon monoxide produced to moles of the oxidizable aromatic compound supplied is preferably in the range of about 0.005: 1 to 0.09: 1, more preferably in the range of about 0.01: 1 to 0.07: 1, even more preferably in the range of about 0.015: 1 to 0.05: 1, and most preferably in the range of 0.02: 1 to 0.4.

The carbon dioxide content in the dry effluent of the oxidation reactor is preferably in the range of about 0.1 to 1.5 mol%, more preferably in the range of about 0.20 to 1.2 mol%, even more preferably in the range of about 0.25 to 0.9 mol% and most preferably in the range from 0.30 to 0.8 mol%. The carbon monoxide content in the dry effluent of the oxidation reactor is preferably in the range of about 0.05 to 0.6 mol%, more preferably in the range of about 0.10 to 0.5 mol%, even more preferably in the range of about 0.15 to 0.35 mol% and most preferably in the range from 0.18 to 0.28 mol%.

Applicants have found that a significant factor to reduce the production of carbon oxides to preferred ranges is to improve the purity of the recycled solvent and to load the oxidizable compound in order to lower the concentration of aromatic compounds without non-aromatic hydrocarbon groups in accordance with the settings of the present invention - this simultaneously reduces the formation of oxides carbon and harmful impurities. Another factor is to improve the distribution of para-xylene and oxidizing agent within the reaction vessel in accordance with the settings of the present invention. Other factors providing the above preferred levels of carbon oxides consist in operating at gradients in the reaction medium, as described in this invention, for pressure, temperature, concentration of the oxidizable compound in the liquid phase and for the oxidizing agent in the gas phase. Other factors providing the above preferred levels of carbon oxides are to be found within the scope of the discoveries in this application for the preferred reaction volumetric rate, pressure, temperature, solvent composition, catalyst composition, and mechanical geometry of the reaction vessel.

One of the possible advantages of working within the preferred ranges of carbon monoxide production is that the use of molecular oxygen can be reduced, although not to stoichiometric values. Despite a good stepwise change in the oxidizing agent and oxidizable compound in accordance with the present invention, the excess oxygen must be maintained above the stoichiometric value that is calculated to load the oxidizable compound separately to allow some loss of carbon oxides and create an excess of molecular oxygen to control the formation of harmful impurities. Specifically, for the case where the charge of the oxidizable compound is xylene, the ratio of molecular oxygen to xylene mass is preferably in the range of about 0.9: 1 to 5: 1, more preferably in the range of about 0.95: 1 to 1.3: 1 and most preferably in the range from 1: 1 to 1.15: 1. Specifically, for xylene loading, the time-averaged molecular oxygen content in the dry effluent of the oxidation reactor is preferably in the range of about 0.1 to 6 mol%, more preferably in the range of about 1 to 2 mol%. and most preferably in the range from 1.5 to 3 mol%.

Another possible advantage of working within the preferred ranges of carbon monoxide production is that less aromatic compounds are converted to carbon oxides and other, less useful forms. This advantage is estimated using the sum of moles of all aromatic compounds leaving the reaction medium divided by the sum of moles of all aromatic compounds entering the reaction medium for a continuous period of time, preferably for one hour, more preferably for one day, and most preferably for 30 consecutive days. This ratio is hereinafter referred to as the “molar survival ratio” for aromatic compounds when passing through the reaction medium and is expressed in numerical percent. If all incoming aromatic compounds leave the reaction medium in the form of aromatic compounds, which are all predominantly in the oxidized form of incoming aromatic compounds, then the molar survival ratio has a maximum value of 100%. If only 1 out of every 100 incoming aromatic molecules is converted to carbon monoxide and / or other non-aromatic molecules (for example, acetic acid) while they pass through the reaction medium, then the molar survival ratio is 99%. Specifically, for the case where the main charge of the oxidizable compound consists of xylene, the molar survival ratio for the aromatic compound when passing through the reaction medium is preferably in the range of about 98 to 99.9%, more preferably in the range of about 98.5 to 99 , 8% and most preferably in the range from 99.0 to 99.7%.

Another aspect of the subject invention involves the production of methyl acetate in a reaction medium containing acetic acid and one or more oxidizable aromatic compounds. Such methyl acetate is relatively volatile compared to water and acetic acid and, therefore, tends to follow the outgoing gas if additional cooling or other work units are not used to separate it and / or to destroy it before releasing the outgoing gas back into the environment. The formation of methyl acetate thus means operating costs as well as capital costs. Apparently, methyl acetate is formed initially by combining the methyl radical, possibly formed as a result of decomposition of acetic acid, with oxygen, to obtain methyl hydroxy peroxide, due to subsequent decomposition with the formation of methanol and, finally, due to the interaction of the obtained methanol with the remaining acetic acid to form methyl acetate. Regardless of the chemical route of formation, as the applicants have established, when the methyl acetate production rate is too low, the formation of carbon oxides is also too low, and the formation of harmful aromatic impurities is too high. If methyl acetate is formed at too high a rate, the formation of carbon oxides is also unnecessarily high, which results in loss of yield over a solvent capable of oxidation, compound and oxidizing agent. When using preferred embodiments of the invention, the ratio of the number of moles of methyl acetate produced to the number of moles of the oxidizable aromatic compound supplied is preferably in the range of about 0.005: 1 to 0.09: 1, more preferably in the range of about 0.01: 1 to 0, 07: 1 and most preferably in the range from 0.02: 1 to 0.04: 1.

When the formation of carbon dioxide, carbon monoxide, their total formation and / or the formation of methyl acetate are below the preferred ranges disclosed in the present invention, or when the molar ratio of the survival of aromatic compounds is above the preferred ranges disclosed in the present invention, the reactivity should be increased or should be lowered SOS. One activity enhancer raises the temperature within the preferred ranges described in the present invention. Another activity enhancer enhances catalytic activity, which is provided by a mixture of catalytic chemicals and a solvent. Typically, increasing concentrations of cobalt and / or bromine will enhance the reactivity if they are used within the preferred ranges. The regulation of the concentration within the reaction medium of other catalytic components and water can also be used to enhance the reaction activity. SOS is reduced by reducing the feed rate of the oxidizable compound and / or by increasing the volume of the reaction medium.

When the formation of carbon dioxide, carbon monoxide, their total formation and / or the formation of methyl acetate are greater than the described preferred ranges, and / or when the molar ratio of the survival of aromatic compounds is higher than the preferred ranges disclosed in the present invention, the preferred control actions are opposite to those given above actions, also at preferred intervals described in the present invention. It was noted that it is especially useful to increase the SOS as much as possible in the described intervals, while maintaining a good oxidation quality, which is estimated by harmful impurities in the STC and in the reaction medium. It was also noted that it is difficult to maintain such a quality of oxidation at such a high SOS and that very great caution is necessary for the following indicators: distribution of raw materials entering the reaction medium; aeration quality throughout the reaction medium; deaeration at the outlet of the reaction medium; SOS oxygen and dissolved oxygen throughout the reaction medium; excess oxidizing agent leaving the reaction medium; the desired spatial gradient of oxygen SOS; the desired spatial concentration gradient of the oxidizable compound; the desired spatial gradient of the concentration of oxidizing agent; head pressure the desired spatial pressure gradient and preferred temperature in the middle of the height of the reaction mixture, and as all discussed in this description. Additionally and in order to achieve lower amounts of carbon dioxide, carbon monoxide and / or their sum and / or to increase the molar survival ratio for aromatic compounds, the applicants have found that it is useful to suppress the concentration of soluble aromatic compounds without non-aromatic hydrocarbon groups within the reaction zone (e.g. isophthalic acid, phthalic acid and benzoic acid); such suppression can be realized by using a cleaner feed of an oxidizable compound and / or a cleaner solvent, especially within the preferred ranges described in the present invention for each component.

In a reaction medium where para-xylene is continuously oxidized to terephthalic acid with the increased SOS described in the present invention, it is preferable that the amount of para-toluic acid in the liquid phase of the reaction medium is maintained in the range of about 200 to 10,000 ppm. more preferably from about 800 to 8000 parts by weight per million, and most preferably from 1600 to 6000 parts by weight per million. In addition, the conversion of para-xylene to terephthalic acid in the reaction medium is preferably maintained at about 50 mol%, more preferably about 90 mol%, even more preferably about 95 mol%. and most preferably over 97 mol%.

In one embodiment of the present invention, it is preferable that one or more of the operating parameters described in the invention (including quantitatively determined by the numerical values of the operating parameters) are held for a significant industrial period of time. Preferably, operation in accordance with one or more of the operating parameters described above is maintained for at least about 1 hour, more preferably at least about 12 hours, even more preferably at least about 36 hours, and most preferably at least 96 hours. Therefore, unless otherwise indicated, the operating parameters described in this case are, as implied, applicable to stationary, optimal / industrial operation — in the absence of start, stop, or partial optimization operations.

It should be noted that for all numerical intervals presented in the present invention, the upper and lower boundaries of the intervals can be independent of each other. For example, a numerical range of 10 to 100 means more than 10 and / or less than 100. Therefore, a range of 10 to 100 provides support for a stated limit of more than 10 (without an upper bound), for a stated limit of less than 100 (without a lower bound ), as well as for the entire interval from 10 to 100 (both with the upper boundary and the lower boundary). In addition, when the definition of “approximately” is used to modify a numerical value, it should be understood that in one embodiment of the invention, the numerical value is an accurate numerical value.

The invention is described in detail using specific references to preferred embodiments thereof, but it should be understood that changes and modifications can be made within the spirit and scope of the present invention.

Claims (30)

1. A method of obtaining a composition of aromatic dicarboxylic acid, including
(a) conducting the oxidation of the multiphase reaction medium in the primary oxidation reactor, resulting in a first suspension;
(b) conducting additional oxidation of at least a portion of said first suspension in a secondary oxidation reactor, wherein said secondary oxidation reactor is a bubble column reactor,
moreover, the method further includes introducing an aromatic compound into said primary oxidation reactor, where at least about 80% by weight of said aromatic compound introduced into said primary oxidation reactor is oxidized in said primary oxidation reactor,
moreover, the head gases are transferred from the top of the secondary oxidation reactor to the primary oxidation reactor.
2. The method according to claim 1, where the specified aromatic compound is a para-xylene.
3. The method according to any one of claims 1 to 2, where step (b) comprises oxidizing the para-toluic acid present in said first suspension.
4. The method according to claim 3, further comprising removing the second suspension from said secondary oxidation reactor, wherein the time-averaged concentration of para-toluic acid in the liquid phase of said second suspension is less than about 50 wt.% Of the time-averaged concentration of para-toluic acid in the liquid phase of the specified first suspension.
5. The method according to claim 4, where the time-averaged concentration of para-toluic acid in the liquid phase of said first suspension is at least about 500 wt. ppm, and where the time-averaged concentration of para-toluic acid in the liquid phase of said second suspension is less than about 250 wt. ppm
6. The method according to claim 1, where the specified primary oxidation reactor is a bubble column reactor.
7. The method according to claim 1, where the specified secondary oxidation reactor is located outside the specified primary oxidation reactor.
8. The method according to claim 7, where at least a portion of the specified secondary oxidation reactor is located along the side of the specified primary oxidation reactor.
9. The method according to claim 1, where the specified secondary oxidation reactor is not a displacement reactor.
10. The method according to claim 1, further comprising withdrawing said first suspension from said primary oxidation reactor through a suspension outlet located between the lower and upper ends of said primary oxidation reactor.
11. The method according to claim 10, where the specified primary oxidation reactor limits it to the primary reaction zone having a maximum height (N p ), where the specified outlet for the suspension is at a distance of at least about 0.1 N p from the bottom and the upper ends of the specified primary reaction zone.
12. The method according to claim 11, where the specified outlet for the suspension is at a distance of at least about 0.25H i from the lower and upper ends of the specified primary reaction zone.
13. The method according to claim 1, where the specified primary oxidation reactor limits the primary reaction zone in it, where the specified secondary oxidation reactor limits the secondary reaction zone in it, where the ratio of the volume of the specified primary reaction zone to the volume of the specified secondary reaction zone is in the range from approximately 4: 1 to 50: 1.
14. The method according to item 13, where the specified primary reaction zone has a ratio of the maximum vertical height to the maximum horizontal diameter in the range of from about 3: 1 to 30: 1, where the specified secondary reaction zone has a ratio of the maximum vertical height to the maximum horizontal diameter in the range from about 1: 1 to 50: 1.
15. The method according to any one of paragraphs.13 and 14, where the ratio of the maximum horizontal diameter of the specified primary reaction zone to the maximum horizontal diameter of the specified secondary reaction zone is in the range from about 0.1: 1 to 0.6: 1, where the ratio of the maximum vertical the height of the specified primary reaction zone to the maximum vertical height of the specified secondary reaction zone is in the range from about 0.1: 1 to 0.9: 1.
16. The method according to item 13, where the specified primary reaction zone has a maximum diameter (D p ), where the volumetric center of gravity of the specified secondary reaction zone is at a horizontal distance of at least about 0.5D p from the volumetric center of gravity of the specified primary reaction zone.
17. The method according to item 13, where the specified primary reaction zone has a maximum height (N p ), where the volumetric center of gravity of the specified secondary reaction zone is located at a vertical distance of less than 0.5N p from the volumetric center of gravity of the specified primary reaction zone.
18. The reactor system for carrying out the method according to claim 1, including
- a primary oxidation reactor restricting the first inlet and the first outlet; and
- a secondary oxidation reactor bounding the second inlet and the second outlet,
where said first outlet is connected by transmitting a fluid stream to said second inlet, wherein said secondary oxidation reactor is a bubble column reactor.
19. The reactor system according to claim 18, wherein said primary oxidation reactor is a bubble column reactor.
20. The reactor system according to any one of paragraphs.18 and 19, where the specified secondary oxidation reactor is located outside the specified primary oxidation reactor.
21. The reactor system according to claim 20, where at least a portion of the specified secondary oxidation reactor is located along the side wall of the specified primary oxidation reactor.
22. The reactor system according to claim 18, wherein said secondary oxidation reactor is not a displacement reactor.
23. The reactor system according to claim 18, wherein said primary oxidation reactor defines a slurry outlet connected by a fluid stream to said secondary oxidation reactor, where said outlet is located between the lower and upper ends of said primary oxidation reactor.
24. The reactor system according to item 23, where the specified primary oxidation reactor limits the primary reaction zone having a maximum height (N p ), where the specified outlet for the suspension is at a distance of at least about 0.1 N p from the bottom and the upper ends of the specified primary reaction zone.
25. The reactor system of claim 24, wherein said outlet for the slurry is at a distance of at least about 0.25H p on the lower and upper ends of said primary reaction zone.
26. The reactor system according to claim 18, wherein said primary oxidation reactor limits a primary reaction zone therein, wherein said secondary oxidation reactor limits a secondary reaction zone therein, wherein a ratio of a volume of said primary reaction zone to a volume of said secondary reaction zone is in the range of about from 4: 1 to 50: 1.
27. The reactor system of claim 26, wherein said primary reaction zone has a ratio of maximum vertical height to maximum horizontal diameter ranging from about 3: 1 to 30: 1, wherein said secondary reaction zone has a ratio of maximum vertical height to maximum horizontal diameter in the range is from about 1: 1 to 50: 1.
28. The reactor system according to any one of paragraphs.26 and 27, where the ratio of the maximum horizontal diameter of the specified zone of primary oxidation to the maximum horizontal diameter of the specified zone of secondary oxidation is in the range from about 0.1: 1 to 0.6: 1, where the ratio of the maximum the vertical height of the specified zone of primary oxidation to the maximum vertical height of the specified zone of secondary oxidation is in the range from about 0.1: 1 to 0.9: 1.
29. The reactor system according to p , where the specified primary reaction zone has a maximum diameter (D p ), where the volumetric center of gravity of the specified secondary reaction zone is at a horizontal distance of at least about 0.5D p from the volumetric center of gravity of the specified primary reaction zone.
30. The reactor system according to p, where the specified primary reaction zone has a maximum height (N p ), where the volumetric center of gravity of the specified secondary reaction zone is located at a vertical distance less than 0.5N p from the volumetric center of gravity of the specified primary reaction zone.
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